Upgrading carbo-metallic oils with used catalyst

ABSTRACT

A process is disclosed for upgrading a hydrocarbon oil feed having a significant content of metals, especially vanadium, to provide a higher grade of oil products by contacting the feed under sorbing conditions in an upgrading zone with a high surface area, high pore volume sorbent material containing an added alkaline metal to neutralize acidic cracking sites. Upgrading conditions are such that coke and metals are deposited on the sorbent in the upgrading zone. Coked sorbent is regenerated by contact with an oxygen containing gas under regeneration conditions to remove the coke, and regenerated sorbent is recycled to the upgrading zone for contact with fresh feed. The added alkaline metal is present on the sorbent in an amount sufficient to neutralize substantially all of the acidic cracking sites and provide a sorbent material having a MAT relative activity in the range of about 0 to about 1 percent. A sorbent composition disclosed comprises a deactivation, spent or equilibrium catalyst withdrawn from an FCC or RCC cracking operation and treated with an alkaline metal additive during the upgrading process and/or prior to use in the upgrading process by impregnation techniques. The alkaline metal additives include water soluble inorganic salts and/or hydrocarbon soluble organo-metallic compounds of select alkaline metals.

TECHNICAL FIELD

This invention is concerned with producing a higher grade petroleum oilhaving lowered contaminating metals and Conradson carbon values from apoor grade of petroleum oil comprising a carbo-metallic oil having highmetals and Conradson carbon values. More particularly, this inventionrelates to the use of solid sorbent particulates for contacting the oilto affect a substantial reduction in the levels of undesired metals andcoke precursors that contribute to losses of selective cracking activityin hydrocarbon conversion catalysts.

BACKGROUND OF THE INVENTION

The introduction of catalytic cracking to the petroleum industry in the1930's constituted a major advance over previous techniques with theobject of increasing the yield of gasoline and its quality. Early fixedbed, moving bed, and fluid bed (FCC) catalytic cracking processesemployed vacuum gas oils (VGO) from crude sources that were consideredsweet and light. The terminology of sweet refers to low sulfur contentand light refers to the amount of material boiling below approximately1,000°-1,025° F.

The catalysts employed in early homogeneous fluid dense beds were of anamorphous siliceous material, prepared synthetically or from naturallyoccurring materials activated by acid leaching. Tremendous strides weremade in the 1950's in FCC technology in the areas of metallurgy,processing equipment, regeneration and new, more-active and more stableamorphous catalysts. However, increasing demand with respect to quantityof gasoline and increased octane number requirements to satisfy the newhigh horsepower-high compression engines being promoted by the autoindustry put extreme pressure on the petroleum industry to increase FCCcapacity and severity of operation.

A major breakthrough in FCC catalysts came in the early 1960's with theintroduction of molecular sieves or zeolites. These materials wereincorporated into the matrix of amorphous and/or amorphous/kaolinmaterials constituting the FCC catalysts of that time. These newzeolitic catalysts, containing a crystalline aluminosilicate zeolite inan amorphous matrix of silica, alumina or silica-alumina, with orwithout kaolin or other clays or the like, were at least 1,000-10,000times more active for cracking hydrocarbons than the earlier amorphouscatalysts containing silica and/or alumina, with or without kaolin orother clays. This introduction of zeolitic cracking catalystsrevolutionized the fluid catalytic cracking process. New innovationswere developed to handle these high activities, such as riser cracking,shortened contact times, new regeneration processes, new improvedzeolitic catalyst developments, and the like.

The new catalyst developments revolved around the development of variouszeolites such as synthetic types X and Y and naturally occurringfaujasites; increased thermal-steam (hydrothermal) stability of zeolitesthrough the inclusion of rare earth ions or ammonium ions viaion-exchanged techniques; and the development of more attritionresistant matrices for supporting the zeolites.

These zeolitic catalyst developments gave the petroleum industry thecapability of greatly increasing throughput of feedstock with increasedconversion and selectivity while employing the units without expansionand without requiring new unit construction.

After the introduction of zeolite containing catalysts, the petroleumindustry began to suffer from a lack of crude availability as toquantity and quality accompanied by increasing demand for gasoline withincreasing octane values. The world crude supply picture changeddramatically in the late 1960's and early 1970's. From a surplus oflight, sweet crudes the supply situation changed to a tighter supplywith an ever increasing amount of heavier crudes with higher sulfurcontents. These heavier and higher sulfur crudes presented processingproblems to the petroleum refiner in that these heavier crudesinvariably also contained much higher metals and Conradson carbonvalues, with accompanying significantly increased asphaltic content.

Fractionation of the total crude to yield cat cracker charge stocks alsorequired much better control to ensure that metals and Conradson carbonvalues were not carried overhead to contaminate the FCC charge stock.The effects of heavy metal and Conradson carbon on a zeolite containingFCC catalyst have been described in the literature as to their highlyunfavorable effect in lowering catalyst activity and selectivity forgasoline production and their equally harmful effect on catalyst life.

As mentioned previously, these heavier crude oils also contained more ofthe heavier fractions and yielded less or lower volume of the highquality FCC charge stocks which normally boil below about 1,025° F. andare usually processed so as to contain total metal levels below 1 ppm,preferably below 0.1 ppm, and Conradson carbon values substantiallybelow 1.0.

With the increasing supply of heavier crudes, which meant lowered yieldsof gasoline, and the increasing demand for liquid transportation fuels,the petroleum industry began a search for processing schemes to utilizethese heavier crudes in producing gasoline. Many of these processingschemes have been described in the literature. These include Gulf'sGulfining and Union Oil's Unifining processes for treating residuum,UOP's Aurabon process, Hydrocarbon Research's H-Oil process, Exxon'sFlexicoking process to produce thermal gasoline and coke, H-Oil'sDynacracking and Phillip's Heavy Oil Cracking (HOC) processes. Theseprocesses utilize thermal cracking or hydrotreating followed by FCC orhydrocracking operations to handle the higher content of metalcontaminates (Ni-V-Fe-Cu-Na) and high Conradson carbon values of 5-15.Some of the drawbacks of these types of processing are as follows:Coking yields thermally cracked gasoline which has a much lower octanevalue than cat cracked gasoline and is unstable due to the production ofgum from diolefins and requires further hydrotreating and reforming toproduce a high octane product; gas oil quality is degraded due tothermal reactions which produce a product containing refractorypolynuclear aromatics and high Conradson carbon levels which are highlyunsuitable for catalytic cracking; and hydrotreating requires expensivehigh pressure hydrogen, multi-reactor systems made of special alloys,costly operations, and a separate costly facility for the production ofhydrogen.

To better understand the reasons why the industry has progressed alongthe processing schemes described, one must understand the known andestablished effects of contaminant heavy metals (Ni-V-Fe-Cu),contaminant alkaline metals (Na) and Conradson carbon on the zeolitecontaining cracking catalysts and the operating parameters of a FCCunit. Heavy metal content and Conradson carbon are two very effectiverestraints on the operation of an FCC unit and may even imposeundesirable restraints on a Reduced Crude Conversion (RCC) unit from thestandpoint of obtaining maximum conversion, selectivity and life.Relatively low levels of these contaminants are highly detrimental to anFCC unit. As heavy metals and Conradson carbon levels are increasedstill further, the operating capcity and efficiency of an RCC unit maybe adversely affected or made uneconomical. These adverse effects occureven though there is enough hydrogen in the feed to produce an idealgasoline consisting of only toluene and isomeric pentenes (assuming acatalyst with such ideal selectivity could be devised).

The effect of increased Conradson carbon is to increase that portion ofthe feedstock converted to coke deposited on the catalyst. In typicalVGO operations employing a zeolite containing catalyst in an FCC unit,the amount of coke deposited on the catalyst averages about 4-5 wt% ofthe feed. This coke production has been attributed to four differentcoking mechanisms, namely, contaminant coke from adverse reactionscaused by metal deposits, catalytic coke caused by acid site cracking,entrained hydrocarbons resulting from pore structure adsorption and/orpoor stripping, and Conradson carbon resulting from pyrolyticdistillation of hydrocarbons in the conversion zone. There has beenpostulated two other sources of coke present in reduced crudes inaddition to the four present in VGO. They are: (1) absorbed and adsorbedhigh boiling hydrocarbons which do not vaporize and cannot be removed bynormally efficient stripping, and (2) high molecular weight nitrogencontaining hydrocarbon compounds adsorbed on the catalyst's acid sites.Both of these two new types of coke producing phenomena add greatly tothe complexity of resid processing. Therefore, in the processing ofhigher boiling fractions, e.g., reduced crudes, residual fractions,topped crude, and the like, the coke production based on feed is thesummation of the four types present in VGO processing (the Conradsoncarbon value generally being much higher than for VGO), plus coke fromthe higher boiling unstrippable hydrocarbons and coke associated withthe high boiling nitrogen containing molecules which are adsorbed on thecatalyst. Coke production on clean catalyst, when processing reducedcrudes, may be estimated as approximately 4-5 wt% of the feed plus theConradson carbon value of the heavy feedstock.

The coked catalyst is brought back to equilibrium activity by burningoff the deactivating coke in a regeneration zone in the presence of air,and the regenerated catalyst is recycled back to the reaction zone. Theheat generated during regeneration is removed by the catalyst andcarried to the reaction zone for vaporization of the feed and to provideheat for the endothermic cracking reaction. The temperature in theregenerator is normally limited because of metallurgical limitations andthe hydrothermal stability of the catalyst.

The hydrothermal stability of the zeolite containing catalyst isdetermined by the temperature and steam partial pressure at which thezeolite begins to rapidly lose its crystalline structure to yield a lowactivity amorphous material. The presence of steam is highly criticaland is generated by the burning of adsorbed and absorbed (sorbed)carbonaceous material which has a significant hydrogen content (hydrogento carbon atomic ratios generally greater than about 0.5). Thiscarbonaceous material is principally the high boiling sorbedhydrocarbons with boiling points as high as 1,500°-1,700° F. or abovethat have a modest hydrogen content and the high boiling nitrogencontaining hydrocarbons, as well as related porphyrins and asphaltenes.The high molecular weight nitrogen compounds usually boil above 1,025°F. and may be either basic or acidic in nature. The basic nitrogencompounds may neutralize acid sites while those that are more acidic maybe attracted to metal sites on the catalyst. The porphyrins andasphaltenes also generally boil above 1,025° F. and may contain elementsother than carbon and hydrogen. As used in this specification, the term"heavy hydrocarbons" includes all carbon and hydrogen containingcompounds that do not boil below about 1,025° F., regardless of whetherother elements are also present in the compound.

As the Conradson carbon value of the feedstock increases, cokeproduction increases and this increased load will raise the regenerationtemperature; thus the unit may be limited as to the amount of feed thatcan be processed because of its Conradson carbon content. Earlier VGOunits operated with the regenerator at 1,150°-1,250° F. A newdevelopment in reduced crude processing, namely, Ashland Oil's "ReducedCrude Conversion Process", as described in the pending U.S. applicationsreferenced below, can operate at regenerator temperatures in the rangeof 1,350°-1,400° F. But even these higher regenerator temperatures placea limit on the Conradson carbon value of the feed at approximately 8,which represents about 12-13 wt% coke on the catalyst based on theweight of feed. This level is controlling unless considerable water isintroduced to further control temperature, which addition is alsopracticed in Ashland's RCC processes.

For purposes of this application, the term "heavy metals" refers tonickel, vanadium, copper and iron, although small amounts of other heavymetal elements may sometimes be present. The metal containing fractionsof crude oils contain these heavy metals in the form of porphyrins andasphaltenes. These metal containing hydrocarbons are deposited on thecatalyst during processing and are cracked in the riser to deposit themetal or are carried over by the coked catalyst as the metallo-porphyrinor asphaltene and converted to the metal oxide during regeneration.Certain of these metals, particularly iron and copper, also may bepresent as the free metal or as inorganic compounds resulting fromeither corrosion of process equipment or contaminants from otherrefining processes. The adverse effects of these metals as taught in theliterature are to cause non-selective or degradative cracking anddehydrogenation to produce increased amounts of coke and light gasessuch as hydrogen, methane and ethane. These mechanisms adversely effectselectivity, resulting in poor yields and quality of gasoline and lightcycle oil. The increased production of light gases, while impairing theyield and selectivity of the processes, also puts an increased demand onthe gas compressor capacity of the refinery. The increase in cokeproduction, in addition to its negative impact on yield, also adverselyaffects catalyst activity and selectivity, greatly increases regeneratorair demand and blower capacity, and may result in uncontrollable and/ordangerous regenerator temperatures.

The heavy metals transfer almost quantitatively from the feedstock oilto the catalyst particles and tend to deposit on interior and exteriorsurfaces of the particles where they can block and/or retard diffusionof the hydrocarbon molecules to the active, i.e., acidic, cracking sitesof the catalyst. In addition, vanadium, and to a lesser extent nickeland the other heavy metals, can migrate to these acidic sites and poisonor kill their catalytic cracking activity. Unless removed by desaltingoperations, as is the usual practice, sodium and other alkali oralkaline earth metals in the crude oil can diffuse to the acidic sitesand also kill their catalytic activity.

These problems of the prior art have been greatly minimized by thedevelopment at Ashland Oil, Inc., of its Reduced Crude Conversion (RCC)Processes described in Ser. No. 094,092 (now U.S. Pat. No. 4,332,673)and the other copending applications referenced below and incorporatedherein by reference. The new process can handle reduced crudes or crudeoils containing high metals and Conradson carbon values previously notsusceptible to direct processing. Normally, these crudes requireexpensive vacuum distillation to isolate suitable feedstocks and produceas a by-product, high sulfur containing vacuum still bottoms. Ashland'sRCC processes avoid all of these prior art disadvantages. However,crudes such as Mexican Mayan or Venezuelan and certain othercarbo-metallic oils and/or oil fractions may contain abnormally highmetal and/or Conradson carbon values. If these poor grades of oil areprocessed in a reduced crude process, they will lead to an uneconomicaloperation because of the high load on the regenerator and/or the highcatalyst replacement rates required to maintain catalyst activity andselectivity. The replacement rate can be as high as 4-8 lbs/bbl which,at today's catalyst prices, can add as much as $2-8/bbl of additionalcatalyst cost to the processing economics. On the other hand, it isdesirable to develop an economical means of processing poor grade oils,such as the Mexican Mayan, because of their availability and cheapnessas compared to Middle East crudes.

The literature suggests many processes for the reduction of metalscontent and Conradson carbon values of reduced crudes and othercontaminated oil fractions. One such process is that described in U.S.Pat. No. 4,263,128 and German Patent No. 29 04 230 assigned to EngelhardMinerals and Chemicals, Inc., which patents are incorporated herein byreference. Basically, these prior art processes involve contacting areduced crude fraction or other contaminated oil with sorbent atelevated temperature in a sorbing zone, such as a fluid bed, to producea product of reduced metal and Conradson carbon value. One of thesorbents described in U.S. Pat. No. 4,263,128 is an inert solidinitially composed of kaolin which has been spray dried to yieldmicrospherical particles having a surface area of about 15 m² /g orless, low porosity, i.e., a pore volume of about 0.20 cc/gm or less, anda catalytic cracking micro-activity (MAT) value not substantiallygreater than 20, and which has been subsequently calcined at hightemperature in an effort to achieve better attrition resistance. As thevanadia content on such sorbents increases into the range of10,000-30,000 ppm, the sorbent begins to have fluidization problemswhich have been overcome previously by removal of most of the spentsorbent inventory and addition of fresh virgin material. This usuallyrequires shutting down the sorbent contacting facility.

DISCLOSURE OF THE INVENTION

The invention is directed to the inclusion of one or more alkaline metaladditives as a select metal oxide or salt into an FCC or RCC catalyst inan amount sufficient to lower its catalytic cracking activity to nearzero. The catalyst is preferably a deactivated, spent or equilibriumcatalyst which has been used previously in an FCC or RCC conversionoperation.

The invention provides a method of producing a high grade of reducedcrude conversion (RCC) feedstocks having lowered metals and Conradsoncarbon values from a poor grade of reduced crude or other carbo-metallicoil having extremely high metals and Conradson carbon values.

The invention may further be used for processing crude oils or crude oilfractions with significant levels of metals and/or Conradson carbon toprovide an improved feedstock for typical fluid catalytic cracking (FCC)processes.

Crude oils or residual fractions from the distillation of crude oils maycontain substantial amounts of heavy metals, such as Ni, V, Fe and Cu,and have high Conradson carbon values. These oils are made suitable forprocessing in a reduced crude conversion (RCC) process or a fluidcatalytic cracking (FCC) process by preliminarily contacting the oilwith a sorbent material exhibiting relatively low or no significantcatalytic cracking activity at elevated temperatures to reduce themetals and Conradson carbon values.

As used throughout this specification, "vanadia" refers collectively tothe oxides of vanadium. It has been found that as vanadium levels buildup on a sorbent, the elevated temperatures encountered in the sorbentregeneration zone cause vanadia, particularly vanadium pentoxide (V₂O₅), to melt and liquid vanadia to flow. The melting and flowing ofvanadia can, particularly at high levels and for sorbent materials withlow surface areas and low pore volumes, coat the outside of sorbentmicrospheres with liquid vanadia and thereby adversely affect sorbentfluidization properties. Any interruption or decrease in particle flowmay result in coalescence between the liquid coated sorbent particles.Once coalescence occurs, fluidization becomes difficult to reinitiate.This results in stoppage of flow in cyclone diplegs, ineffectiveoperation of cyclones, rapid increases in the loss of the sorbent, andmay finally result in unit shutdown.

According to the present invention, the adverse effects of vanadium aregreatly reduced by contacting contaminated oil feeds with a sorbenthaving a relatively high pore volume and surface area to immobilizevanadium oxides by adsorbing deposited vanadia within the porousstructure of the sorbent particles during feed pretreatment. The selectsorbent of the invention is preferably a deactivated, spent orequilibrium (used) catalyst which has been treated with alkaline metalsalts to neutralize residual acid cracking sites and yield a materialwhich esssentially no catalytic cracking activity. Significant crackingactivity is preferably avoided in the upgrading process to maximize theyield of high quality liquid product for use as feedstock in asubsequent FCC or RCC conversion process employing a highly selectivecracking catalyst under optimum cracking conditions. This combination ofupgrading and conversion processes substantially increases the yields ofhigh octane liquid fuels from poor grades of crude or resid oils.

The method of addition of the alkaline metal additives to decrease thecatalytic cracking activity of used FCC or RCC catalysts can be byaddition during the upgrading process cycle or by impregnation of thecatalyst particles, such as spray dried microspheres, prior to use inthe process cycle. The resulting sorbent may possess less crackingactivity than the previously mentioned and utilized sorbents and has ahigher pore volume and a higher internal surface area to adsorb vanadia.Internal adsorption of vanadia deposited on the sorbent duringprocessing of the oil for metal and/or Conradson carbon removal greatlyreduces the tendency for particle coalescence at high vanadia loadings.

The invention thus provides an improved sorbent and an improved methodfor treatment of petroleum oil feeds containing significant levels ofvanadium (at least about 1.0 ppm). More particularly, improved particlepore structure is provided in the sorbent to reduce particle coalescenceand loss of fluidization caused by vanadium contaminants in oil feeds ofall types utilized in FCC and/or RCC operations. Some crude oils andsome FCC charge stocks from the distillation of crude oils containsignificant amounts (greater than 1.0 ppm) of heavy metals. They mayalso contain significant amounts (greater than 1.0 ppm) of alkalinemetals, such as sodium, magnesium and calcium. Residual fractions fromcrude oil distillation have even greater amounts of heavy metals andalkaline metals and may also have high Conradson carbon values. Theinvention is particularly useful in the pretreatment of such residualfractions to provide carbometallic oil feeds for RCC units.

It is to be understood that the catalyst particles can be of any size,depending on the size appropriate to the upgrading process in which thesorbent is to be employed. Thus, while a fluidizable size is preferred,larger particles of the select sorbent may be employed, such as in amoving bed for upgrading unvaporized feeds. If the used catalyst is toolarge for the sorbent process selected, it may be ground up or otherwisecomminuted to provide a sorbent of the desired particle size.

The problems of the prior art caused by vanadium containing feedstocksare overcome by employing the select sorbent of this invention. Thisinvention is especially effective in the treatment of reduced crudes andother carbo-metallic feeds with high metals, high vanadium to nickelratios and high Conradson carbon values. Such MRS (Metals RemovalSystem) feeds with high metal and Conradson carbon values are preferablycontacted in a riser with the inert solid sorbent of increased porevolume and very low catalytic cracking activity at temperatures aboveabout 900° F. Residence time of the oil in the riser is below 5 seconds,preferably below 3 seconds and most preferably in the range of 0.5 to 2seconds. The preferred sorbent is a spray dried catalyst composition inthe form of microspherical particles generally in the size range of 10to 200 microns, preferably 20 to 150 microns and more preferably between40 and 80 microns, to ensure adequate fluidization properties. Theseparticles preferably are comprised of recovered microspheres ofequilibrium catalysts from FCC and/or RCC conversion operations.

The MRS feed is introduced at the bottom of the riser and contacts thesorbent at a temperature of 1,150°-1,400° F. to yield a temperature atthe exit of the riser in the sorbent disengagement vessel ofapproximately 900°-1,000° F. Along with the MRS feed, water, steam,naphtha, flue gas, or other vapors or gases may be introduced to aid invaporization of the feed and act as a lift gas for providing the shortresidence time desired.

Coked sorbent is abruptly separated from the hydrocarbon vapors at theexit of the riser by a ballistic separation apparatus employing thevented riser concept developed by Ashland Oil, Inc., and described inU.S. Pat. Nos. 4,066,533 and 4,070,159 to Myers, et al., which patentsare incorporated herein by reference.

During the course of the treatment in the riser, the metal and Conradsoncarbon compounds are deposited on the sorbent. After ballisticseparation from the product vapors, the coked sorbent is deposited as adense but fluffed bed at the bottom of the disengagement vessel,transferred to a stripper and then to a regeneration zone in a"combustor" vessel. In the combustor, the coked sorbent is contactedwith an oxygen containing gas to remove the coke deposits by combustionof this carbonaceous material to form carbon oxides and yield aregenerated sorbent containing less than 0.5 wt% carbon based on weightof sorbent, preferably less than 0.2 wt% carbon and more preferably lessthan 0.10 wt% carbon. The regenerated sorbent is then recycled to thebottom of the riser where it again joins high metal and Conradson carboncontaining feed to repeat the cycle.

At the elevated temperatures encountered in the regeneration zone, thevanadium deposited on the sorbent in the riser is converted to vanadiumoxides, in particular, vanadium pentoxide. The melting point of vanadiumpentoxide is much lower than the temperatures encountered in theregeneration zone. Thus, it can become a mobile liquid and flow acrossthe sorbent surface so as to cause pore plugging and particlecoalescence, which phenomena are particularly prevalent with sorbents ofthe type described in the literature.

This application describes a new approach to offsetting the adverseeffects of vanadium pentoxide by employing high surface area, high porevolume catalytic materials whose catalytic cracking activity has beenneutralized by the incorporation of one or more alkaline metals or theiroxides, salts or other compounds. This alkaline metal additive may beintroduced into the catalytic matererial by conventional impregnationtechniques, or during the upgrading of a carbo-metallic oil byintroducing the additive at one or more select points in the upgradingsystem to affect neutralization of the active catalytic cracking sites.

The preferred alkaline metal additives for neutralizing the acidiccracking sites include the following metals, their oxides, salts, and/ororgano-metallic compounds: Li, Na, K, Rb, Cs, Mg, Ca, Sr, and Ba. Otheradditives which may be used include the other elements of Groups IA andIIA of the Periodic Chart of Elements and their compounds. Thesealkaline metal additives may be impregnated on the catalyst in amountssufficient to yield a concentration of the alkaline metal element on thesorbent in the range of about 0.1 to 10 percent, more preferably about0.2 to 5 percent and most preferably about 0.5 to 2 percent by weight.As an alternative, the amount of additive metal used may be based on itsneutralization equivalency and the concentration of residual acidicsites on the catalyst that are to be neutralized as explained in moredetail below. If added instead during the treatment process, the metalelements may build up to these concentrations on equilibrium sorbent andbe maintained at these levels by controlling the rate of additiveaddition relative to the rate of sorbent replacement.

The sorbents of this invention are prepared from solids of highactivity, such as deactivated, spent and/or equilibrium catalystswithdrawn from FCC and/or RCC processing operations. The catalystswithdrawn from FCC or RCC operations contain catalytically activealuminosilicate zeolite in a matrix which may itself posses significantcracking activity or have no significant cracking activity, e.g.,kaolin. The catalytically active cracking sites present in the matrixand in the aluminosilicate zeolite are neutralized by the selectalkaline metal additives previously described to yield a sorbentpossessing extremely low or essentially no catalytic activity and highadsorptive powers for liquid vanadia. These sorbents preferably have asurface area above about 20 m² /g, more preferably above about 50 m² /gand most preferably above about 70 m² /g; a pore volume greater thanabout 0.2 cc/g, more preferably greater than about 0.3 cc/gm and mostpreferably greater than about 0.4 cc/gm; and a microactivity value, asmeasured by ASTM Test Method No. D3907-80 and reported as MAT relativeactivity where relative activity of the standard is 100%, of below about1%, more preferably below about 0.6%, still more preferably in the rangeof about 0 to 0.2% and most preferably below about 0.05%. These relativeactivities correspond approximately to respective MAT values in volumepercent conversion of below about 35%, more preferably below about 30%,still more preferably in the range of about 0 to about 20% and mostpreferably below about 7.0%.

The select sorbents and the upgrading processes described herein arepreferably employed to provide an RCC feedstock for the carbo-metallicoil conversion (RCC) processes described more fully below and in thecopending patent applications referenced below (all pending patentapplications referred to in this specification are assigned or are to beassigned to Ashland Oil, Inc.). Although the term "RCC" is anabbreviation derived from a heavy oil conversion process known asReduced Crude Conversion, this term is employed herein to refergenerally to those processes for cracking and/or reforming oil feedshaving levels of heavy metals, alkaline metals and/or Conradson carbongenerally higher than those usually employed in Fluid Catalytic Cracking(FCC) operations. The term "MRS" is employed herein to refer generallyto those processes for upgrading a poor grade of feedstock oil bycontacting this oil with a sorbent material. The invention furthercontemplates combined MRS and RCC processes wherein feedstock for theRCC process is comprised of an upgraded liquid product from the MRSprocess and the sorbent for the MRS process is comprised of an RCCequilibrium catalyst wthdrawn from the RCC process and treated with theselect alkaline metal additives of the invention. Similarly, theinvention contemplates combined MRS and FCC processes wherein feedstockfor the FCC process is comprised of an upgraded liquid product from theMRS process and the sorbent for the MRS process is comprised of an FCCequilibrium catalyst withdrawn from the FCC process and treated withselect alkaline metal additives of the invention. It is alsocontemplated that the MRS sorbent may be comprised of combinations ofRCC and/or FCC catalysts from partially or wholly integrated conversionoperations and/or from entirely independent conversion operations, andthat one fraction of the MRS product may be used as an RCC feedstockwhile another fraction is being used simultaneously as an FCC feedstock.

BRIEF DESCRIPTION OF THE DRAWINGS

The invention may be further understood by reference to the descriptionof the best and other illustrative modes for carrying out the inventiontaken in conjunction with the accompanying drawings in which:

FIG. 1 is a graph illustrating the numerical relationship between MATrelative activity and MAT volume percent conversion.

FIG. 2 is a schematic diagram of an apparatus for carrying out anupgrading process according to the invention.

FIG. 3 is a schematic block diagram illustrating one flow arrangement ofa MRS process for upgrading poor grades of reduced crude integrated witha RCC process for catalytically cracking the MRS liquid product and forproviding used cracking catalyst to be neutralized and used as sorbentin the MRS process.

FIG. 4 is a schematic diagram of a contactor and combustor apparatus forthe metals removal system (MRS) of the combined MRS-RCC complex of FIG.3.

FIG. 5 is a schematic diagram of a reactor and regenerator apparatus forthe reduced crude conversion (RCC) system of the combined MRS-RCCcomplex of FIG. 3.

FIG. 6 is a graph illustrating the changes in particle coalescenceproperties with increasing amounts of vanadium both on untreated claysorbent and on used catalyst sorbent treated with an alkaline metaladditive according to the invention.

FIG. 7 is a graph showing the change in MAT relative activity of sorbentas the alkaline metal concentration on sorbent is increased.

BEST AND OTHER ILLUSTRATIVE MODES FOR CARRYING OUT THE INVENTION

A catalytically active cracking catalyst containing zeolite as apromoter is first manufactured and then utilized in an FCC or RCCprocess over a period of time to yield a less active or "used"equilibrium catalyst. Thereafter, the residual active cracking sites ofthe used catalyst are neutralized with an alkaline metal additive toyield the select sorbent of this invention and the select sorbent isemployed to upgrade a metals and Conradson carbon containing oil for useas feed in the same or a different FCC or RCC process.

The catalytic material, which preferably is but need not be adeactivated, spent or equilibrium catalyst from an FCC or RCC process,is treated with a sufficient amount of the alkaline metal additive toneutralize substantially all of the acidic catalytic cracking sites andyield a sorbent with a micro-activity test (MAT) value of preferablyless than about 30 volume percent, more preferably less than about 20volume percent, still more preferably less than about 10 volume percentand most preferably about 8 volume percent or less as measured by ASTMTest Method No. D3907-80. Thorough neutralization of the acid catalyticcracking sites avoids significant cracking activity in the MRS upgradingprocess and maximizes the yield of high quality liquid product availablefor use as feed in a subsequent FCC or RCC conversion operationemploying a highly selective cracking catalyst under optimum crackingconditions. This optimized upgrading followed by optimized conversionsubstantially increases the overall yields of high octane liquid fuelsfrom poor grades of crude or resid containing oils.

It is not proposed to define the exact mechanism by which the sorbent ofthis invention reduces the adverse effects of vanadia depositsaccumulating on the sorbent material. However, the sorbent of thisinvention enjoys a higher surface area and pore volume than thatexhibited or possessed by conventional sorbents described in theliterature. Without intending to be bound by any one theory orhypothesis, it is believed that this increased surface area and porevolume permits large amounts of vanadia to be adsorbed within the porestructure before vanadia can accumulate on exterior particle surfaces inamounts sufficient to cause pore plugging and/or particle coalescence.The invention contemplates the lower oxidation states of vanadium aswell as vanadium pentoxide. Furthermore, in treating a sulfur containingfeed and regenerating sorbent in the presence of an oxygen containinggas, the vanadium may also be present as a sulfide, sulfate and/oroxysulfide. The methods of making and using the select sorbents of theinvention are described fully in the following sections of thisspecification.

CATALYST COMPOSITIONS

The select MRS sorbents of this invention are prepared from solidshaving some catalytic activity such as zeolites in a matrix of clays,kaolin, silica, alumina, smectites and other 2-layered lamellarsilicates, silica-alumina, or a combination of two or more of thesematrix materials. The surface area of these sorbents are preferablyabove 50 m² /g and preferably they have a pore volume substantially inexcess of 0.2 cc/g.

A preferred matrix material for catalyst compositions to be usedultimately as a MRS sorbent is a semi-synthetic combination of clay andsilica-alumina as described in U.S. Pat. No. 3,034,994, the entiredisclosure of which is incorporated herein by reference. Preferably theclay is mostly a kaolinite and is combined with a syntheticsilica-alumina hydrogel or hydrosol. This synthetic component formspreferably about 15 to 75 percent, more preferably about 20 to 25percent, of the final catalyst composition by weight. The proportion ofclay is such that the catalyst preferably contains after forming, about10 to 75 percent, more preferably about 30 to 50 percent, clay based ontotal weight of catalyst, including zeolite and/or other promoters. Themost preferred composition of the matrix contains approximately twice asmuch clay as synthetically derived silica-alumina. The syntheticallyderived silica-alumina may contain 55 to 95 percent by weight of silica(SiO₂), preferably 65 to 85 percent, most preferably about 75 percent.Catalysts wherein the gel matrix consists entirely of silica gel oralumina gel are also included.

Various processes may be used in preparing the synthetic silica-aluminamatrix, such as those described in U.S. Pat. No. 3,034,994. A preferredone of these processes involves gelling an alkali metal silicate with aninorganic acid while maintaining the pH on the alkaline side. An aqueoussolution of an acidic aluminum salt is then intimately mixed with thesilica hydrogel so that the aluminum salt solution fills the silicahydrogel pores. The aluminum is thereafter precipitated as a hydrousalumina by the addition of an alkaline compound, the hydrous aluminacombining with silica at the surface of the silica hydrogen pores. Thehydrous gel is then processed, for instance, by separating a part of thewater on vacuum filters and then drying, or more preferably, by spraydrying the hydrous gel slurry to produce microspheres. The dried productis washed to remove sodium and sulfate ions, either with water or a veryweak acid solution. The resulting product is then dried to a lowmoisture content, usually less than 25 percent by weight, e.g., 10percent to 20 percent by weight, to provide a finished catalyst product.

The silica-alumina hydrogel slurry may be filtered and washed in gelform to affect purification of the gel by the removal of dissolvedsalts. This may enhance the formation of a continuous phase in the spraydried microspherical particles. If the slurry is prefiltered and washedand it is desired to spray dry the filter cake, the latter may bereslurried with enough water to produce a pumpable mixture for spraydrying. The spray dried product may then be washed again and given afinal drying in the manner previously described.

Prior to final drying of the matrix material, it is preferably mixedwith a highly active catalytic promoter, such as a zeolite. The zeolitecomponent is preferably a synthetic faujasite which possesses silica toalumina ratios in the range from about 2.5 to 7.0, preferably 3.0 to 6.0and most preferably 4.5 to 6.0. Synthetic faujasites are widely knowncrystalline aluminosilicate zeolites and common examples are the x and Ytypes of zeolites commercially available from the Davison Division of W.R. Grace and Company and the Linde Division of Union CarbideCorporation. The ultra-stable hydrogen exchanged zeolites, such asZ-14XS and Z-14US from Davison, are also particularly suitable. Inaddition to faujasites, other preferred types of zeolitic materials aremordenite and erionite.

The most commonly used Y-type faujasites may be prepared as described inU.S. Pat. Nos. 3,130,007 and 4,010,116, the entire disclosures of whichare incorporated herein by reference. The aluminosilicates of the latterpatent have high silica (SiO₂) to alumina (Al₂ O₃) molar ratios,preferably above 4, to give high thermal stability.

Zeolites used for catalytic cracking are usually made in the sodium formand then exchanged with polyvalent cations to reduce the Na₂ O contentto less than about 1.0 percent by weight, and preferably less than 0.1percent by weight. Procedures for removing alkali metals and puttingzeolites in the proper form are well-known in the art and are described,for example, in U.S. Pat. Nos. 3,293,192; 3,402,996; 3,446,727;3,449,070; and 3,537,816; the entire disclosures of said patents beingincorporated herein by reference.

The amount of zeolitic material dispersed in the matrix material basedon the final composite product should be at least about 5 weightpercent, preferably in the range of about 10 to 50 weight percent, mostpreferably about 20 to 40 weight percent.

Crystalline aluminosilicate zeolites exhibit acidic sites on bothinterior and exterior surfaces, with the largest proportion of totalsurface area and cracking sites being located internal to the zeolitecrystals within crystalline micropores. Zeolites are usuallycrystallized as regularly shaped discrete particles of approximately 0.1to 10 microns in size and, accordingly, this is the size range normallyprovided by commercial catalyst suppliers. The particle size of thezeolites are preferably in the lower portion of this size range in orderto increase exterior (portal) surface areas. The preferred zeolites arethermally stabilized with hydrogen and/or rare earth ions and are steamstable to about 1,650° F.

The zeolites can be suitably dispersed in the matrix materials for useas cracking catalysts by methods well-known in the art, such as thosedisclosed, for example, in U.S. Pat. Nos. 3,140,249 and 3,140,253 toPlank, et al.; U.S. Pat. No. 3,660,274 to Blazek, et al.; U.S. Pat. No.4,010,116 to Secor, et al.; U.S. Pat. No. 3,944,482 to Mitchell, et al.;and U.S. Pat. No. 4,079,019 to Scherzer, et al.; the entire disclosuresof said patents being incorporated herein by reference.

After introduction of the zeolite, the composition is preferablyslurried and spray dried to form microspheres of matrix material havingzeolite particles (crystals) dispersed uniformly therein. As analternative to intimately mixing a preformed zeolite with a slurriedmatrix material, the zeolite may be formed in situ by silication of aclay component of the matrix as described in U.S. Pat. No. 4,010,116.The finished catalyst should contain from 5 to 50% by weight zeolite,preferably a rare earth or ammonia exchanged sieve of either or both theX and Y variety. To further enhance catalyst hydrothermal stability, thespray dried microspheres containing a previously exchanged sieve arepreferably calcined and further exchanged with rare earth or ammonia tocreate an exceptionally stable sieve.

Commercial zeolite-containing catalysts are available with carrierscontaining a variety of metal oxides and combination thereof, includingfor example silica, alumina, magnesia and mixtures thereof, and mixturesof such oxides with clays as described, for example, in U.S. Pat. No.3,034,948, the entire disclosure of which is incorporated herein byreference. One may for example select any of the zeolite-containing FCCcatalysts which are suitable for production of gasoline from vacuum gasoils. However, certain advantages may be attained by judicious selectionof catalysts having marked resistance to metals, e.g., RCC molecularsieve catalysts. A metal resistant zeolite catalyst is, for instance,described in U.S. Pat. No. 3,944,482, which catalyst contains 1 to 40weight percent of a rare earth-exchanged zeolite, the balance being arefractory metal oxide matrix having a specified pore volume and sizedistribution. Other catalysts described as "metals-tolerant" aredisclosed in an article by Cimbalo, et al., entitled "Deposited MetalsPoison FCC Catalysts", Oil and Gas Journal, May 15, 1972, pp. 112-122,the entire contents of which are incorporated herein by reference.

In progressive flow FCC and RCC reactors and/or in progressive flow MRScontactors, it is preferred to employ solids having an overall particlesize in the range of about 10 to about 200 microns, more preferablyabout 20 to 150 microns and most preferably about 40 to about 80microns. A useful catalyst sorbent may have a skeletal density of about150 pounds per cubic foot and an average particle size of about 60-70microns, with less than 10% of the particles having a size less thanabout 40 microns and less than 80% having a size less than about 50-60microns.

Although a wide variety of other catalysts, including bothzeolite-containing and non-zeolite-containing, may be employed in thepractice of the invention, the following are examples of commerciallyavailable catalysts which may be employed in practicing the invention:

                  TABLE 1                                                         ______________________________________                                        Spe-       Weight Percent                                                     cific      Zeo-                                                               Sur-       lite                                                               face       Con-                                                               m.sup.2 /g tent    Al.sub.2 O.sub.3                                                                      SiO.sub.2                                                                          Na.sub.2 O                                                                          Fe.sub.2 O                                                                          TiO.sub.2                         ______________________________________                                        AGZ-290 300    11.0    29.5  59.0 0.40  0.11  0.59                            GRZ-1   162    14.0    23.4  69.0 0.10  0.4   0.9                             CCZ-220 129    11.0    34.6  60.0 0.60  0.57  1.9                             Super DX                                                                              155    13.0    31.0  65.0 0.80  0.57  1.6                             F-87    240    10.0    44.0  50.0 0.80  0.70  1.6                             FOX-90  240    8.0     44.0  52.0 0.65  0.65  1.1                             HFZ-20  310    20.0    59.0  40.0 0.47  0.54  2.75                            HEZ-55  210    19.0    59.0  35.2 0.60  0.60  2.5                             ______________________________________                                    

The AGZ-290, GRZ-1, CCZ-220 and Super DX catalysts referred to above areproducts of W. R. Grace and Company. F-87 and FOX-90 are products ofFiltrol, while HFZ-20 and HEZ-55 are products of Engelhard/Houdry. Theabove are properties of virgin catalyst and, except in the case ofzeolite content, are adjusted to a water-free basis, i.e., based onmaterial ignited at 1,750° F. The zeolite content is derived bycomparison of the X-ray intensities of a catalyst sample and of astandard material composed of high purity sodium Y zeolite in accordancewith draft #6, dated Jan. 9, 1978, of proposed ASTM Standard Methodentitled "Determination of the Faujasite Content of a Catalyst".

Among the above-mentioned commercially available catalysts, the Super Dfamily and especially a catalyst designated GRZ-1 are particularlypreferred. For example, Super DX has given particularly good resultswith relatively high metals. The GRZ-1, although substantially moreexpensive than the Super DX at present, appears somewhat moremetals-tolerant.

Although not yet commercially available, it is believed that the bestcatalysts for carrying out the present invention are those which arecharacterized by matrices with feeder pores having large minimumdiameters and large mouths to facilitate diffusion of high molecularweight molecules through the matrix to the portal surface area ofmolecular sieve particles within the matrix. Such matrices preferablyalso have a relatively large pore volume in order to soak up unvaporizedportions of a carbo-metallic oil feed. Thus, significant numbers ofliquid hydrocarbon molecules can diffuse to active catalytic sites bothin the matrix and in sieve particles at the surface of the matrix. Ingeneral it is preferred to employ catalysts having a total pore volumegreater than 0.2 cc/gm, preferably at least 0.4 cc/gm, more preferablyat least 0.6 cc/gm and most preferably in the range of 0.7 to 1.0 cc/gm,and with matrices wherein at least 0.1 cc/gm, and preferably at least0.2 cc/gm, of said total pore volume is comprised of feeder pores havingdiameters in the range of about 400 to about 6,000 angstrom units, morepreferably in the range of about 1,000 to about 6,000 angstrom units.These catalysts and the method for making the same are described morefully in copending international application Ser. No. PCT/US81/00492filed in the U.S. Receiving Office on Apr. 10, 1981, in the names ofAshland Oil, Inc., et al., and entitled "Large Pore Catalysts for HeavyHydrocarbon Conversion", the entire disclosure of said application beingincorporated herein by reference.

The catalyst composition to be used as sorbent may include variousadditives for passivating the non-selective catalytic activity of heavymetal deposits on the conversion catalyst. The catalyst may beimpregnated with these additives during manufacture and/or the additivesmay be introduced into the riser, the regenerator or other conversionsystem components during the FCC or RCC conversion process in which thecatalyst is initially employed. Preferred additives for this purposeincludes those disclosed in U.S. patent application Ser. No. 263,395,filed on May 13, 1981, in the name of William P. Hettinger, Jr., andentitled "Passivating Heavy Metals in Carbo-Metallic Oil Conversion",the entire disclosure of said U.S. application being incorporated hereinby reference.

Other metal additives specifically for reducing the adverse effects ofvanadium on catalyst may be employed as described in PCT InternationalApplication Ser. No. PCT/US81/00356 filed in the U.S. Receiving Officeon Mar. 19, 1981, in the names of Ashland Oil, Inc., et al., andentitled "Immobilization of Vanadia Deposited on Catalytic MaterialsDuring Carbo-Metallic Oil Conversion", the entire disclosure of said PCTInternational Application being incorporated herein by reference. Thecatalyst composition may further include special ingredients fortrapping vanadium within the porous matrix of the particles as describedin U.S. patent application Ser. No. 252,967 filed on Apr. 10, 1981, inthe names of James D. Carruthers, et al., and entitled "Trapping ofMetals Deposited on Catalytic Materials During Carbo-Metallic OilConversion", the entire disclosure of which is incorporated herein byreference.

PREPARATION OF SORBENT

Prior to inclusion of one or more of the alkaline metal additives of thepresent invention into a catalyst of the foregoing type, the catalyst ispreferably used in a FCC conversion operation of the progressive flowtype or in one of Ashland's RCC conversion operations. In suchoperations, the appropriate FCC or RCC feed is introduced into thebottom of a riser reactor of the type illustrated in FIG. 5 along with aheated suspension of the catalyst. Steam, naphtha, water, flue gasand/or some other diluent is preferably also introduced into the riserto aid in dispersing and mixing the feed and catalyst and as a source ofvapor for accelerating the feed and catalyst to achieve the vaporvelocities and residence times desired. As the feed travels up theriser, it is catalytically cracked to form basically five products knownin the industry as dry gas, wet gas, cat naphtha, light cycle oil, heavycycle oil and/or slurry oil. At the upper end of the riser, the catalystparticles are ballistically separated from product vapors using thevented riser techniques described elsewhere in this specification. Thecatalyst, which then contains the coke and heavy metal deposits formedin the riser, is sent to the regenerator to burn off the coke and theseparated product vapors are sent to a fractionator for furtherseparation and treatment to provide the five basic products indicated.

The preferred riser conditions of an FCC process providing a source ofequilibrium FCC catalyst for practicing the invention are summarized inTable 2. The preferred riser conditions of Ashland's RCC processesproviding a source of equilibrium RCC catalyst for practicing theinvention are summarized in Table 3. In these tables, the abbreviationsused have the following meanings: "Temp." for temperature, "Dil." fordiluent, "pp" for partial pressure, "wgt" for weight, "V" for vapor,"Res." for residence, "C/O" for catalyst to oil ratio, "Cat." forcatalyst, "bbl" for barrel, "MAT" for micro-activity by the MAT testusing a standard Davison feedstock, "Vel." for velocity, "cge" forcharge, "d" for density and "Reg." for regenerated.

                  TABLE 2                                                         ______________________________________                                        FCC RISER CONDITIONS                                                                        Board                                                                         Operating     Preferred                                         Parameter     Range         Range                                             ______________________________________                                        Feed Temp.     400-800° F.                                                                          400-650° F.                               Steam Temp.    200-500° F.                                                                          300-400° F.                               Reg. Catalyst Temp.                                                                         1000-1400° F.                                                                        1175-1350° F.                              Riser Exit Temp.                                                                             900-1200° F.                                                                         925-1050° F.                              Pressure        0-100 psia   10-50 psia                                       Water/Feed    0.01-0.15     0.01-0.10                                         Dil. pp/Feed pp                                                                             0.15-2.0      0.25-1.0                                          Dil. wgt/Feed wgt                                                                           ≦0.2   0.01-0.1                                          V. Res. Time  0.1-5 sec.    0.5-3 sec.                                        C/O, wgt.     4-12          5-10                                              Lbs. Cat./bbl Feed                                                                          0.01-2.0      0.05-1                                            Inlet Cat. MAT                                                                              >60 vol. %    70-85 vol. %                                      Outlet Cat. MAT                                                                             ≧55 vol. %                                                                           ≧65 vol. %                                 V. Vel.       25-90 ft./sec.                                                                              30-60 ft./sec.                                    V. Vel./Cat. Vel.                                                                           ≧1.0   1.2-2.0                                           Dil. Cge. Vel.                                                                              5-90 ft./sec. 10-50 ft./sec.                                    Oil Cge. Vel. 1-50 ft./sec.  5-50 ft./sec.                                    Inlet cat. d. 1-9 lbs./ft..sup.3                                                                           2-6 lbs./ft..sup.3                               Outlet cat. d.                                                                              1-6 lbs./ft..sup.3                                                                           1- 3 lbs./ft..sup.3                              ______________________________________                                    

                  TABLE 3                                                         ______________________________________                                        RCC RISER CONDITIONS                                                                         Board                                                                         Operating    Preferred                                         Parameter      Range        Range                                             ______________________________________                                        Feed Temp.      400-800° F.                                                                         400-650° F.                               Steam Temp.     200-500° F.                                                                         300-400° F.                               Reg. Catalyst Temp.                                                                          1100-1500° F.                                                                       1275-1450° F.                              Riser Exit Temp.                                                                              900-1400° F.                                                                        950-1100° F.                              Pressure         0-100 psia  10-50 psia                                       Water/Feed     0.05-0.30    0.05-0.15                                         Dil. pp/Feed pp                                                                              0.25-3.0     1.0-2.5                                           Dil. wgt/Feed wgt                                                                            ≦0.4  0.1-0.3                                           V. Res. Time   0.1-5 sec.   0.5-3 sec.                                        C/O, wgt.      3-18         5-12                                              Lbs. Cat./bbl Feed                                                                           0.1-4.0      0.2-2.0                                           Inlet Cat. MAT >50 vol. %   ≧60 vol. %                                 Outlet Cat. MAT                                                                              ≧20 vol. %                                                                          ≧40 vol. %                                 V. Vel.        ≧25 ft./sec.                                                                        ≧30 ft./sec.                               V. Vel./Cat. Vel.                                                                            ≧1.0  1.2-2.0                                           Dil. Cge. Vel. 5-90 ft./sec.                                                                              10-50 ft./sec.                                    Oil Cge. Vel.  1-50 ft./sec.                                                                               5-50 ft./sec.                                    Inlet cat. d.  1-9 lbs./ft..sup.3                                                                          2-6 lbs/ft..sup.3                                Outlet cat. d. 1-6 lbs./ft..sup.3                                                                          1- 3 lbs/ft..sup.3                               ______________________________________                                    

The gas for burning coke off of the catalyst in the regenerator may beany gas which can provide oxygen to convert carbon to carbon oxides. Airis highly suitable for this purpose in view of its ready availability.The amount of air required per pound of coke for combustion depends uponthe desired carbon dioxide to carbon monoxide ratio in the effluentgases and upon the amount of other combustible materials present in thecoke, such as hydrogen, sulfur, nitrogen and other elements capable offorming gaseous oxides at regenerator conditions. Following cokeburnoff, the regenerated catalyst is recycled to the riser for contactwith fresh feed.

At such time that the metal level on the catalyst becomes intolerablyhigh such that catalyst activity and/or selectivity declines tounacceptable levels, virgin catalyst is added and deactivatedequilibrium catalyst is withdrawn at one or more addition-withdrawallocations in the conversion/regeneration cycle. Equilibrium catalyst ispreferably withdrawn from the FCC or RCC process when its MAT crackingactivity has decreased to a value less than about 60 volume percent,more preferably less than about 55 volume percent, which valuescorrespond to MAT relative activities of about 20 percent and 10percent, respectively.

The withdrawn catalyst is then treated with the select alkaline metalsalts described below to neutralize residual acid cracking sites andyield a material with essentially no catalytic cracking activity. Themethod of addition of the alkaline metal additives to the used FCC orRCC catalyst can be by addition during the upgrading MRS process cycleof by impregnation of the catalyst particles prior to use in the MRSprocess cycle.

The additive alkaline metals of the invention include the metal elementsof Groups IA and IIA of the Periodic Chart Of The Elements. Theinvention also contemplates treatment of the catalyst with otheradditive metals that possess alkaline or amphoteric properties, such aszinc and cadmium in Group IIB, gallium and indium in Group IIIA,germanium and tin in Group IVA and arsenic, antimony and bismuth inGroup VA. The preferred alkaline metals are Li, Na, K, Rb, Cs, Mg, Ca,Sr and Ba. The more preferred alkaline metal additives are compounds ofsodium, potassium, magnesium, calcium or barium or a mixture of thecompounds of these metals.

Where the metal additive is introduced directly into the upgradingprocess, that is into the riser, into the regenerator or into anintermediate component, the alkaline metal additives are preferablyorgano-metallic compounds soluble in the hydrocarbon feed or in ahydrocarbon solvent miscible with the feed, or are water soluble saltsof sufficient solubility for the quantities required to be added towater streams fed to the upgrading process. Various inorganic compoundsare also soluble in hydrocarbon solvents. The invention therefore is notlimied to the examples given.

The organo-metallic additives would include alcoholates, esters,phenolates, naphthenates, carboxylates, dienyl sandwich compounds. Theorgano-metallic additives are preferably introduced directly into thehydrocarbon upgrading zone, preferably near the bottom of the riser, sothat the alkaline metal additive will be deposited on the sorbent. Whenthe alkaline metal of the invention reaches the regenerator, its oxideis formed either by decomposition of the additive directly to the metaloxide or by decomposition of the additive to the free metal which isthen oxidized under the regenerator conditions. This provides intimatecontact between the alkaline metal additive and the acidic crackingsites and is believed to be an effective means for neutralizing theacidic cracking sites present in the matrix and zeolite structures.

The alkaline metal additive is introduced into the riser by mixing itwith the feed in an amount sufficient to give at least about a 1:1ratio, preferably a ratio in the range of about 1.2 to about 2.0, ofalkaline neutralization equivalents to acid sites still remaining on thepartially deactivated catalyst (alkaline newutralization ratio), theneutralization equivalency of an alkaline metal being depending upon itsvalence in the additive compound. For example, sodium has aneutralization equivalency of 1 and magnesium has a neutralizationequivalency of 2. Therefore, a 1:1 alkaline neutralization ratio for agiven number of acid sites requires an equal number of sodium atoms butonly one-half that number of magnesium atoms. A 1:1 alkalineneutralization ratio for alkaline metals having valences greater than 1assumes that the acidic sites are close enough together for a singlemetal atom to neutralize multiple acidic sites. Since this assumptionmay not hold near the end of neutralization where residual acidic sitesmay occur at significant intervals apart, alkaline neutralization ratiostoward the upper end of the specified range of 1.2 to 2.0 are preferredfor alkaline metals having valence states greater than 1.

The residual acidity of withdrawn FCC or RCC equilibrium catalyst may bedetermined by techniques well known in the art such as, for example, bytitration with a standard basic solution of known volume and strength.One such technique is described in an article entitled "Technique forMeasuring Surface Acidity of Catalyst", by H. A. Benesi in the Journalof Physical Chemistry, Vol. 61, page 970, 1957, and involves titrationof the catalyst with n-butyl amine in a non-aqueous solvent such asbenzene. Other techniques for determining the number of acid sites perweight of catalyst are described in a multi-volume treatise entitledAdvances in Catalysis, 1978, in the section by H. A. Benesi commencingat page 97 of vol. 27. The entire disclosures of said article andsection by H. A. Benesi are incorporated herein by reference.

If the alkaline metal additive is added directly to the deactivated,spent or equilibrium catalyst at some time before the sorbent isintroduced into the upgrading system, the alkaline metal additive ispreferably a water soluble organic and/or inorganic salt of thespecified metals, such as the acetate, halide, nitrate, sulfate, sulfiteand/or carbonate. These water soluble alkaline metal additives may alsobe introduced into the treatment process along with water containingstreams, such as used to cool the regenerator or to lift, fluidize orstrip the sorbent.

After neutralization of the catalytic acid sites, the select sorbents ofthe invention have a micro-activity (MAT) value in volume percentconversion as measured by ASTM Test Method No. D-3907-80 preferablybelow about 35%, more preferably below about 20%, still more preferablybelow about 10% and most preferably in the range of about 0% to about7%. In general, it is preferred to employ a sorbent having essentiallyno cracking activity so as to provide very low levels of conversion atlow residence times. Catalytic cracking activity is preferably minimizedduring the upgrading process to provide the maximum amount of a highquality liquid feed which is low in contaminates and can therefore becracked later either in a downstream conversion zone or in a separateconversion process operated at optimum conditions for selectivecatalytic cracking. Catalytic cracking during the upgrading process,which necessarily takes place in the presence of relatively high levelsof contaminates, is much more non-selective and produces both loweroctane products and higher yields of coke and light gases than catalyticcracking of better quality feed in the absence of such adversecontaminate levels.

The cracking capabilities of the sorbent may be expressed in terms ofthe conversion produced during actual operation of the process and/or interms of conversion produced in standard catalyst activity tests. Forexample, the preferred sorbent may be defined as a catalyst which, inits neutralized state, exhibits a specified activity expressed as apercentage in terms of MAT (microactivity test) conversion. For purposesof the present invention, the foregoing percentage is the volumepercentage of standard feedstock which the sorbent under evaluation willconvert to 430° F. end point gasoline, lighter products and coke at 900°F., 16 WHSV (weight hourly space velocity calculated on a moisture freebasis using clean catalyst which has been dried at 1100° F., weighed andthen conditioned for a period of at least 8 hours at about 25° C. and50% relative humidity, until about one hour or less prior to contactingthe feed), and 3 S/O (sorbent to oil weight ratio) by ASTM D-32 MAT testD-3907-80. The standard feedstock is an appropriate high grade petroleumoil, e.g., a sweet light primary gas oil, such as that used by theDavison Division of W. R. Grace and defined as follows:

    ______________________________________                                        API Gravity at 60° F., degrees                                                               31.9                                                    Specific Gravity at 60° F., g/cc                                                             0.8708                                                  Ramsbottom Carbon, wt. %                                                                            0.09                                                    Conradson carbon, wt. % (est.)                                                                      0.04                                                    Carbon, wt. %         84.92                                                   Hydrogen, wt. %       12.94                                                   Sulfur, wt. %         0.68                                                    Nitrogen, ppm         305                                                     Viscosity at 100° F., centistokes                                                            10.36                                                   Watson K Factor       11.93                                                   Aniline Point         182                                                     Bromine No.           2.2                                                     Paraffins, Vol. %     31.7                                                    Olefins, Vol. %       1.8                                                     Naphthenes, Vol. %    44.0                                                    Aromatics, Vol. %     22.0                                                    Average Molecular Weight                                                                            284                                                     Nickel                Trace                                                   Vanadium              Trace                                                   Iron                  Trace                                                   Sodium                Trace                                                   Chlorides             Trace                                                   B S & W               Trace                                                   Distillation          ASTM D-1160                                             IBP                   445                                                     10%                   601                                                     30%                   664                                                     50%                   701                                                     70%                   734                                                     90%                   787                                                     FBP                   834                                                     ______________________________________                                    

The gasoline end point and boiling temperature-volume percentrelationships of the products produced in the MAT conversion test mayfor example be determined by simulated distillation techniques, forexample by modification of the gas chromatographic "Sim-D" technique ofASTM D-2887-73. The results of such simulations are in reasonableagreement with the results obtained by subjecting larger samples ofmaterial to standard laboratory distillation techniques. Conversion iscalculated by subtracting from 100 the volume percent (based on freshfeed) of those products heavier than gasoline which remain in therecovered product.

On pages 935-937 of Hougen and Watson, "Chemical Process Principles",John Wiley & Sons, Inc., N.Y. (1947), the concept of "Activity Factors"is discussed. This concept leads to the use of "relative activity" tocompare the cracking activity of an operating sorbent against a standardmaterial (catalyst). Relative activity measurements facilitaterecognition of how the quantity requirements of various materials differfrom one another. Thus, relative activity is a ratio obtained bydividing the weight of a standard or reference material (catalyst) whichis or would be required to produce a given level of conversion by theweight of an operating material (whether proposed or actually used)which is or would be required to produce the same level of conversionusing the same or equivalent feedstock under the same or equivalentconditions. Said ratio of material weights may be expressed as anumerical ratio, but preferably is converted to a percentage basis. Thestandard material is preferably chosen from among catalyst, such as forexample, zeolite fluid cracking catalysts, and is chosen for its abilityto produce a predetermined level of conversion using a standard feedunder the conditions of temperature, WHSV, catalyst to oil ratio andother conditions set forth in the preceding description of the MATconversion test and in ASTM D-32 MAT test D-3907-80. Conversion is thevolume percentage of feedstock that is converted to 430° F. endpointgasoline, lighter products and coke. For standard feed, one may employthe above-mentioned light primary gas oil, or equivalent.

For purposes of conducting relative activity determinations, one mayprepare a "standard catalyst curve", a chart or graph of conversion (asabove defined) vs. reciprocal WHSV for the standard catalyst andfeedstock. A sufficient number of runs is made under ASTM D-3907-80conditions (as modified above) using standard feedstock at varyinglevels of WHSV to prepare an accurate "curve" of conversion vs. WHSV forthe standard feedstock. This curve should traverse all or substantiallyall of the various levels of conversion, including the range ofconversion within which it is expected that the operating sorbent willbe tested. From this curve, one may establish a standard WHSV for testcomparisons and a standard value of reciprocal WHSV corresponding tothat level of conversion which has been chosen to represent 100%relative activity in the standard catalyst. For purposes of the presentdisclosure, the aforementioned reciprocal WHSV and level of conversionare, respectively, 0.0625 and 75%. In testing an operating sorbent ofunknown relative activity, one conducts a sufficient number of runs withthat sorbent under D-3907-80 conditions (as modified above) to establishthe level of conversion which is or would be produced with the operatingsorbent at standard reciprocal WHSV.

Then, using the above-mentioned standard catalyst curve, one establishesa hypothetical reciprocal WHSV constituting the reciprocal WHSV whichwould have been required, using the standard catalyst, to obtain thesame level of conversion which was or would be exhibited by theoperating sorbent at standard WHSV. The relative activity may then becalculated by dividing the hypothetical reciprocal WHSV of the standardcatalyst by the actual reciprocal WHSV of the test sorbent. The resultis relative activity expressed in terms of a decimal fraction, which maythen be multiplied by 100 to convert to % relative activity (relativeactivity may also be expressed as follows: relative activity at constantconversion is equal to the ratio of the WHSV of the test sorbent dividedby the WHSV of the standard catalyst). To simplify this calculation, aMAT conversion vs. relative activity curve was developed utilizing astandard catalyst of 75 vol.% conversion to represent 100% relativeactivity. One such curve is shown in FIG. 1. In applying the results ofthis determination, a relative activity of 0.01, or 1%, means that itwould take 100 times the amount of the operating sorbent to give thesame conversion as the standard catalyst, i.e., the production sorbentis 1% as active as the reference catalyst.

The alkaline metal is added to neutralize the active acidic crackingsites remaining in the zeolite and/or matrix of preferably adeactivated, spent or equilibrium catalyst. FIG. 7 illustrates theamount of sodium required to neutralize the active cracking sites inthree different commercially available equilibrium catalysts. Thecircles represent a catalyst containing 40 percent zeolite in the virginstate and designated here as "Catalyst D", the diamonds represent acatalyst with 60 percent of the zeolite content of Catalyst D, and thesquares a catalyst containing 50 percent of the zeolite content ofCatalyst D.

The neutralization effectiveness of sodium was evaluated in thelaboratory by depositing sodium on a virgin catalyst. The virgincatalyst was subjected to severe thermal and steaming conditionsaccording to a test sequence designated as calcining, impregnation andsteaming (CIS) in order to approximate an equilibrium catalystimpregnated with sodium and used in an upgrading process. According tothe CIS test, fresh catalyst is calcined at 1200° F. for 3 hours in ashallow bed, 100 gms of the dried material is then vacuum impregnatedwith about 0.5, 1.0 and 2.0 wt% of added sodium. Either aqueoussolutions of sodium sulfate or pentane solutions of metal organiccomplexes are employed. Excess solvent is removed at 0.1 mm Hg pressure.The impregnated catalyst is oxidized at 1000° F. for 3 hours using ashallow bed and muffle furnace. After oxidation, the oxidized materialis steamed at 1450° F. for 5 hours according to the Steaming TortureTest of Table 4. Samples are then tested for MAT activity,selectivities, surface area, zeolite and metal concentrations.

                  TABLE 4                                                         ______________________________________                                        STEAMING TORTURE TEST                                                         FOR ACCELERATED DEACTIVATION                                                  OF FLUID CRACKING CATALYST                                                    PURPOSE:                                                                      This method outlines the deactivation procedure                               for impregnated and oxidized catalyst by hydro-                               thermal treatment before the catalytic cracking                               activity is determined in the Micro-Activity                                  Test (MAT).                                                                   TEST PARAMETERS:                                                              Fluid-bed, quartz reactor, diameter-2.5 cm. ID                                Catalyst load - 75 grams                                                      Heat-up rate - 3° C./min.                                              Nitrogen gas velocity - 0.31 cm/sec. at 788° C.                        Steam gas velocity - 10.9 cm/sec at 788° C.                            Steam rate - 97% gas                                                          TEST PROCEDURE:                                                               Weigh fresh catalyst.                                                         Charge loaded reactor to furnace at room temperature.                         Begin flow of nitrogen at 0.05 SCFH rate.                                     Heat the reactor at maximum rate and begin the steaming                       period when 15° C. of desired steaming temperature is                  reached.                                                                      Start a flow of 100% steam at this temperature. Steam                         flow is continued for 5 hours. A nitrogen flow is                             used in addition to steam to provide constant fluidization.                   Hold reactor at constant desired steaming temperature                         of 788° C. for duration of steaming.                                   After 5 hours, stop the steam and nitrogen flow to                            the reactor.                                                                  Remove reactor from furnace and allow to cool in air                          to ambient conditions.                                                        Submit samples for testing.                                                   CATALYST ANALYSES:                                                            Deactivated catalyst is analyzed for the following parameters:                Surface Area by BET Method                                                    MAT by Micro Activity Test                                                    Mercury Pore Volume                                                           Zeolite, Percent Relative Intensity to Na-Y by                                X-Ray diffraction.                                                            ______________________________________                                    

As shown in FIG. 7, approximately 10,000 ppm of added sodium wasrequired to neutralize the acid cracking sites of the "Catalyst D"series of samples so as to yield a sorbent with essentially no (zero)MAT activity. In further tests, a virgin catalyst containing about 9 wt%zeolite and 0.70 wt% sodium and having a MAT relative activity of about100% after steaming was impregnated with sufficient sodium to raise thesodium content to 1.05 wt% and lower the MAT relative activity to about0.35%, and another virgin catalyst containing about 18 wt% zeolite and0.15% sodium and having a MAT relative activity of about 310% aftersteaming was impregnated with sufficient sodium to raise the sodiumcontent to 1.65 wt% and lower the MAT relative activity about 0.06%,which equates to essentially no catalytic activity. (While 0.06%relative activity represents a finite but low level of conversion, i.e.,8 vol.%, this conversion of the feed is attributed to thermal as opposedto catalytic cracking.)

These series of tests should not be considered as establishing anylimits on the amounts of alkaline metal required to neutralize the acidcracking sites of a deactivated, spent or equilibrium FCC or RCCcatalyst. They are indicative of the results that one can obtain bycarefully controlling the addition of the alkaline metal. The exactamount of alkaline metal required in an operating situation is dependentupon the severity of prior operations that the FCC or RCC equilibriumcatalyst has experienced, such as metal deposition, hydrothermalhistory, regeneration temperatures, original zeolite content, stabilityand type of zeolite and the like, and upon the parameters of thecatalyst treatment (neutralization) process.

UPGRADING PROCESSES

While the processes described in Ashland's RCC applications can handlecrude oils, reduced crudes and other petroleum fractions containing highmetals and Conradson carbon values not susceptible previously to directprocessing, certain crudes such as Mexican Mayan or Venezuelan andcertain other types of oil feeds contain abnormally high heavy metalsand Conradson carbon values. If these very poor grades of oil areprocessed in a RCC process, they may lead to uneconomical operationsbecause of high heat loads on the regenerator and/or high catalystreplacement rates to maintain adequate catalyst activity and/orselectivity. In order to improve the quality of such very poor grades ofoil, such as those containing more than 50 ppm heavy metals and/or morethan 8 weight percent Conradson carbon and preferably more than 100 ppmheavy metals and/or more than 10 weight percent Conradson carbon, theseoils are pretreated with the select sorbent of the invention to reducethe levels of these contaminants to the aforementioned or lower values.One upgrading process thay may be utilized for this purpose is describedin U.S. Pat. No. 4,263,128 issued Apr. 21, 1981, in the name of David B.Bartholic, the entire disclosure of said patent being incorporatedherein by reference. A preferred mode of upgrading residual oils inpracticing the present invention is disclosed in PCT patent applicationSer. No. PCT/US81/00648, filed May 13, 1981, in the names of Oliver J.Zandona, Dwight F. Barger, Paul W. Walters and Lloyd E. Busch andentitled A COMBINATION PROCESS FOR UPGRADING RESIDUAL OILS, the entiredisclosure of said PCT application being incorporated herein byreference.

Certain feeds processed in the past by RCC units because of their metalsand/or Conradson carbon content could be processed instead in FCC unitsif the levels of these contaminates could be lowered by an amountsufficient to yield an FCC feed. For example, a relatively small amount(5-25%) of reduced crude or other heavy hydrocarbon feedstock may bemixed with VGO and the blend upgraded to provide an FCC feed.Alternatively, a reduced crude or other heavy fraction may first beupgraded and then a relatively small amount (5-25%) of the upgradedfraction mixed with VGO to provide an FCC feed. Therefore the presentinvention also contemplates upgrading RCC type feeds to a qualitycapable of being cracked economically in conventional FCC processes andapparatuses.

Representative feedstocks contemplated for use with the inventioninclude whole crude oils; light fractions of crude oils such as lightgas oils (LGO), heavy gas oils (HGO), and vacuum gas oils (VGO); andheavy fractions of crude oils such as topped crude, reduced crude,vacuum fractionator bottoms, other fractions containing heavy residua,coal-derived oils, shale oils, waxes, untreated or deasphalted residua,and blends of such fractions with gas oils and the like. A poor grade ofFCC feedstock for which the invention is particularly useful is onehaving a Conradson carbon value greater than about 1.0, preferablygreater than about 2.0, and containing greater than 0.5 ppm, preferablygreater than 1.0 ppm, total heavy metals. A poor grade of RCC feedstockis one having a Conradson carbon value greater than about 8, preferablygreater than about 10, and containing greater than 50 ppm, preferablygreater than 100 ppm, total heavy metals.

The greater the heavy metals content of the particular type (e.g. RCC orFCC) of oil stock to be upgraded and the greater the proportion ofvanadium in that heavy metal content, the more advantageous the sorbentand processes of this invention becomes. A high vanadium feed for FCCprocessing is one having more than about 0.1 ppm vanadium, morepreferably more than about 1.0 ppm and most preferably about 2.0 toabout 5.0 ppm. A high vanadium feed for RCC processing is one havingmore than about 1.0 ppm vanadium, preferably more than about 5.0 ppm andmost preferably more than about 20 ppm. In either case, the invention isparticularly effective where the weight ratio of vanadium to nickel inthe feed is in the range of from about 1:3 to 5:1, preferably greaterthan about 1:1, more preferably greater than about 2:1 and mostpreferably greater than about 3:1.

Although it is necessary to maintain low sodium values in the feed topractice the art of catalytic cracking as taught in the literature, itis not necessary to maintain low sodium levels in the feed to practicethe present invention. The invention is particularly useful in treatingfeeds containing greater than 50 ppm sodium, preferably greater than 100ppm and more preferably greater than 200 ppm.

A particularly preferred FCC feedstock for upgrading by the process ofthe invention includes VGO, LGO and/or HGO mixed with 5 to 25 wt% of areduced crude to provide a blend having a total heavy metals contentgreater than about 4 ppm of which at least about 2 ppm is vanadium, avanadium to nickel atomic ratio of at least about 1.0, and a Conradsoncarbon value of at least about 2.0. A particularly preferred RCCfeedstock for upgrading by the process of the invention includes areduced crude comprising 70% or more of a 650° F.+ material having afraction greater than 20% boiling above about 1,025° F. at atmosphericpressure, a total heavy metals content of greater than about 50 ppm ofwhich at least about 25 ppm is vanadium, a vanadium to nickel atomicratio of at least about 1.0, and a Conradson carbon value greater thanabout 8.0. This RCC feed may also have a hydrogen to carbon ratio ofless than about 1.8 and coke precursors in an amount sufficient to yieldgreater than about 10% coke by weight based on fresh feed.

The process according to the methods of the invention will produce cokein amounts of 1 to 14 percent by weight based on weight of fresh feed.This coke is laid down on the sorbent in amounts in the range of about0.3 to 3 percent by weight of sorbent, depending upon the sorbent to oilratio (weight of sorbent to weight of feedstock) in the riser and uponthe Conradson carbon value of the particular feed. The severity of theprocess should be sufficiently low so that conversion of the feed togasoline and lighter products is below 20 volume percent, preferablybelow 10 volume percent and more preferably below 5 volume percent. Evenat these low levels of severity, the upgrading process is effective toreduce Conradson carbon values by at least 20 percent, preferably in therange of 40 to 80 percent, and heavy metals content by at least 50percent, preferably in the range of 75 to 90 percent.

The invention may be practiced in a wide variety of apparatus. However,the preferred apparatus includes means for rapidly vaporizing as muchfeed as possible and efficiently admixing feed and sorbent (although notnecessarily in that order), for causing the resultant mixture to flow asa dilute suspension in a progressive-flow mode, and for separating thesorbent from the upgraded feed and lighter products at the end of apredetermined residence time or times, it being preferred that all or atleast a substantial portion of the product should be abruptly separatedfrom at least a portion of the sorbent.

For example, the apparatus may include, along its elongated reactionchamber, one or more points for introduction of carbo-metallic feed, oneor more points for introduction of diluent material, one or more pointsfor withdrawal of products and one or more points for withdrawal ofsorbent.

It is preferred that the contact chamber, or at least the major portionthereof, be more nearly vertical than horizontal and have alength-to-diameter ratio of at least about 10, more preferably about 20or 25 or more. Use of a vertical riser is preferred. This MRS riser isreferred to in this specification as the "contactor" to avoid confusionwith the FCC or RCC catalytic conversion riser. If tubular, thecontactor can be of uniform diameter throughout or may be provided witha continuous or step-wise increase in diameter along the contact path tomaintain or vary the velocity along the flow path.

The means for introducing feed, sorbent and other materials into thecontact chamber may range from open pipes to sophisticated jets or spraynozzles, it being preferred to use means capable of breaking up oratomizing the liquid feed into fine droplets. A particularly preferredatomization and solids contacting process for heavy oils is disclosed inU.S. patent application Ser. No. 263,391, filed on May 13, 1981, in thename of William P. Hettinger, Jr., et al., and entitled "Process forCracking High-Boiling Hydrocarbons Using High Pore Volume, Low DensityCatalyst", the entire disclosure of which is incorporated herein byreference. Preferably, the sorbent, liquid water (when used) and freshfeed are brought together in an apparatus similar to that disclosed inU.S. patent application Ser. No. 969,601 of George D. Myers, et al.,filed Dec. 14, 1978, the entire disclosure of which is incorporatedherein by reference. According to a particularly preferred embodimentbased on a suggestion which is believed to have emanated from Mr.Stephen M. Kovach at Ashland, the liquid water and carbo-metallic oil,prior to their introduction into the riser, are caused to pass through apropeller, apertured disc, or any other appropriate high shear agitatingmeans for forming a "homogenized mixture" containing finely-divideddroplets of oil and/or water with the droplets of one preset in acontinuous phase of the other.

In general, the charging means (for sorbent, feed and diluent) and thecontactor configuration are such as to provide a relatively highvelocity of flow and a relatively dilute suspension of sorbent and feed.For example, the vapor and sorbent velocity in the contactor will beusually at least about 25 or more, typically at least about 35 feet persecond. This velocity is preferably at least about 55, more preferablyabout 75 and most preferably about 100 feet per second and higher. Thevapor velocity at the top of the contactor may be lower than that at thebottom with the above specified velocities being at the bottom. Thevelocity capabilities of the contactor will in general be sufficient toprevent substantial build-up of a sorbent bed in the bottom or otherportions of the contactor, whereby the sorbent loading in the contactorcan be maintained below about 3 or 4 pounds, as for example about 1.2pounds, and below about 2 pounds, as for example 0.8 pounds, per cubicfoot, respectively, at the upstream (e.g., bottom) and downstream (e.g.,top) ends of the contactor.

The progressive flow mode involves, for example, flowing of sorbent,feed, diluent and products as a stream in a positively controlled andmaintained direction established by the elongated nature of the reactionzone. This is not to suggest however that there must be strictly linearflow. Turbulent flow and "slippage" of sorbent relative to vapor mayoccur to some extent especially in certain ranges of vapor velocity andsorbent loading. It is preferable to employ sufficiently low sorbentloadings and sufficiently high vapor velocities to restrict slippage andback-mixing and provide very short contact times.

Most preferably the contactor is one which abruptly separates asubstantial portion or all of the vaporized products from the sorbent atone or more points along a riser, and preferably separates substantiallyall of the vaporized products from the sorbent at the downstream end ofa riser. A preferred type of contactor embodies ballistic separation ofthe sorbent and products; that is, sorbent is projected in a directionestablished by the riser tube, and is caused to continue its motion inthe general direction so established, while the products, having lessermomentum, are caused to make an abrupt change of direction, resulting inan abrupt, substantially instantaneous separation of product fromsorbent. In a preferred embodiment referred to as a vented riser, theriser tube is provided with a substantially unobstructed dischargeopening at its downstream end for discharge of sorbent. An exit portadjacent the downstream end of the riser tube receives the products. Thedischarge opening communicates with a sorbent flow path which includes astripper and sorbent regenerator, while the exit port communicates witha product flow path which is substantially or entirely separated fromthe sorbent flow path and leads to separation means for separating theproducts from the relatively small portion of sorbent, if any, whichmanages to gain entry to the product exit port. Examples of ballisticseparation apparatuses and techniques as above described are found inU.S. Pat. Nos. 4,066,533 and 4,070,159 to Myers, et al., the entiredisclosures of which are incorporated herein by reference.

The mode of sorbent/product separation presently deemed best forpracticing the present invention is disclosed in a U.S. patentapplication Ser. No. 263,394, filed on May 13, 1981, in the names ofPaul W. Walters, Roger M. Benslay, and Dwight F. Barger, entitledCARBO-METALLIC OIL CONVERSION WITH BALLISTIC SEPARATION. The ballisticseparation step preferably includes at least a partial reversal ofdirection by the product vapors upon discharge from the riser tube; thatis, the product vapors make a turn or change of direction which mayexceed 90° at the riser tube outlet. This may be accomplished forexample by providing an annular cup-like member surrounding the risertube at its upper end, the ratio of cross-sectional area of the annulusof the cup-like member relative to the cross-sectional area of the risertube outlet being preferably in the range of about 0.7 to about 1.2 andmore preferably about 1.0. Preferably the lip of the cup is slightlyupstream of, or below the downstream end or top of the riser tube. Bymeans of a product vapor line communicating with the interior of the cupbut not the interior of the riser tube and having its inlet positionedwithin the cup interior upstream of the riser tube outlet, productvapors emanating from the riser tube and entering the cup by an abruptchange of direction are transported away from the cup to auxiliarysorbent and product separation equipment downstream of the cup, e.g., acyclone separator. Such an arrangement can produce a high degree ofcompletion of the separation of sorbent from product vapors at thevented contactor outlet, so that the required amount of auxiliarysorbent separation equipment such as cyclones is greatly reduced, withconsequent large savings in capital investment and operating cost.

Preferred conditions for operation of the process are described below.Among these are feed, sorbent and contact temperatures, contact and feedpressures, residence times and levels of conversion.

In conventional FCC operations with VGO, the feedstock is customarilypreheated, often to temperatures significantly higher than are requiredto make the feed sufficiently fluid for pumping and for introductioninto the riser. For example, preheat temperatures as high as about 700°or 800° F. have been reported. In the MRS process as presentlypracticed, it is preferred to restrict preheating of the feed, so thatthe feed is capable of absorbing a larger amount of heat from thesorbent while the sorbent raises the feed to a viscosity breaking(visbreaking) temperature, at the same time minimizing the amount offeedstock cracked or otherwise converted to gaseous products. Thus,where the nature of the feedstock permits, it may be fed at ambienttemperature. Heavier stocks may be fed at preheat temperatures of up toabout 600° F., typically about 200° F. to about 500° F., but higherpreheat temperatures are not necessarily excluded.

The larger the amount of coke which must be burned from a given weightof sorbent, the greater the potential for exposing the feed to excessivetemperatures. In addition, the large pore structure and pore volume ofthe select sorbents may be susceptible to thermal and/or hydrothermaldegradation at temperatures near the upper end of the temperature rangefor sorbent regeneration. The use of such large pore sorbents forupgrading carbo-metallic feeds creates a need for regenerationtechniques which will not destroy the sorbent matrix by exposure tohighly severe temperatures and steaming. Such need can be met by amulti-stage regeneration process which includes conveying coked sorbentinto a first regeneration zone and introducing oxidizing gas thereto.The amount of oxidizing gas that enters said first zone and theconcentration of oxygen or oxygen-bearing gas therein are sufficient foronly partially affecting the desired conversion of coke on the sorbentto carbon oxide gases. The partially regenerated sorbent is then removedfrom the first regeneration zone and is conveyed to a secondregeneration zone. Oxidizing gas is introduced into the secondregeneration zone to provide a higher concentration of oxygen oroxygen-containing gas than in the first zone, to complete the removal ofcarbon to the desired level. The regenerated sorbent may then be removedfrom the second zone and recycled to the contactor for contact withfresh feed. An example of such a multi-stage regeneration process isdescribed in U.S. patent application Ser. No. 969,602 in the name ofGeorge D. Myers, et al., filed Dec. 14, 1978, the entire disclosure ofwhich is incorporated herein by reference. Another example may be foundin U.S. Pat. No. 2,398,739.

Multi-stage regeneration offers the possibility of combiningoxygen-deficient regeneration with the control of the CO:CO₂ molarratio. Thus, about 50% or more, more preferably about 65% to about 95%,and more preferably about 80% to about 95% by weight of the coke on thesorbent immediately prior to a final regeneration stage may be removedin one or more preceding stages of regeneration in which the molar ratioof CO₂ :CO is maintained at a level substantially below 5, preferablyabout 4 or less, more preferably about 2 or less and most preferablyabout 1.0 or less. In combination with the foregoing, the last 5% ormore, or 10% or more by weight of the coke originally present, up to theentire amount of coke remaining after the preceding stage or stages, canbe removed in a subsequent stage of regeneration in which more oxygen ispresent. Such process is susceptible of operation in such a manner thatthe total flue gas recovered from the entire, completed regenerationoperation contains little or no excess oxygen, e.g., on the order ofabout 0.2 mole percent or less, or as low as about 0.1 mole percent orless, which is substantially less than the 2 mole percent which has beensuggested for catalysts. Thus, multi-stage regeneration is particularlybeneficial in that it provides another convenient technique forrestricting regeneration heat transmitted to fresh feed via regeneratedsorbent and/or for reducing the potential for thermal degradation of thesorbent, while simultaneously affording an opportunity to reduce thecarbon level on regenerated sorbent to those very low percentages (e.g.,about 0.1% or less) which particularly enhance the upgrading capacity ofthe sorbent. For example, a two-stage regeneration process may becarried out with the first stage burning about 80% of the coke at a bedtemperature of about 1400° F. to produce CO and CO₂ in a molar ratio ofCO₂ /CO of about 1 and the second stage burning about 20% of the coke ata bed temperature of about 1370° F. to produce substantially all CO₂mixed with free oxygen. Use of the gases from the second stage ascombustion-supporting gases for the first stage, along with additionalair introduced into the first stage bed, may result in an overall CO₂ toCO ratio of about 1.5, with a sorbent residence time of about 5 to 20minutes or more in the two zones. Moreover, where the regenerationconditions are substantially more severe in the first zone than in thesecond zone (e.g., higher zone or localized temperatures and/or moresevere steaming conditions), that part of the regeneration sequencewhich involves the most severe conditions is performed while there isstill an appreciable amount of coke on the sorbent. Such operation mayprovide some protection of the sorbent from the more severe conditions.A particularly preferred embodiment of the invention is two-stagefluidized regeneration at a maximum temperature of about 1500° F. with areduced temperature by as much as about 20° to 30° F. or more in thedense phase of the second stage as compared to the dense phase of thefirst stage, and with reduction of carbon on sorbent to about 0.1% orless by weight of sorbent in the second zone.

The CO₂ /CO ratio also may be decreased by providing chlorine in anoxidizing atmosphere within the regeneration zone, the concentration ofchlorine preferably being in the range of about 100 to about 400 ppm.This method of decreasing the CO₂ /CO ratio is disclosed in copendingapplications Ser. No. 247,751 filed Mar. 23, 1981, for "Addition ofMgCl₂ to Catalyst" and Ser. No. 246,782 filed Mar. 23, 1981, for"Addition of Chlorine to Regenerator", both in the name of George D.Myers, the entire disclosures of said copending applications beingincorporated herein by reference.

U.S. patent applications Ser. Nos. 94,092 and 94,227 (now U.S. Pat. Nos.4,332,673 and 4,354,923 respectively) referenced below disclosureprocesses for the conversion of carbo-metallic oils to liquid fuel inwhich various regeneration techniques are employed that assist incontrolling the heat load in the regeneration step. Another method ofcontrolling the heat load in catalyst regeneration is disclosed in U.S.patent application Ser. No. 251,032 for "Addition of Water toRegenerator Air" filed Apr. 3, 1981, by George D. Myers, et al., theentire disclosure of this application being incorporated herein byreference. These or similar methods may also be employed for controllingthe heat load in regenerating the sorbent of the present invention.

The regeneration apparatus for removing coke deposited on the sorbent bythe MRS upgrading process is referred to in this specification as a"combustor" in order to avoid confusing this apparatus with FCC and RCCcatalyst regenerators, although these three apparatuses may be of thesame or a similar construction. In most circumstances, it will beimportant to ensure that no adsorbed oxygen-containing gases are carriedinto the contactor by sorbent recycled from the combustor. Thus,whenever such action is considered necessary, the sorbent dischargedfrom the combustor may be stripped with appropriate stripping gases toremove oxygen containing gases. Such stripping may for instance beconducted at relatively high temperatures, for example about 1350° F. toabout 1370° F., using steam, nitrogen or other inert gas as thestripping gas(es). The use of nitrogen and other inert gases isbeneficial from the standpoint of avoiding a tendency towardhydrothermal sorbent degradation which may result from the use of steam.

With respect to the tolerance levels of heavy metals on the sorbentitself, such metals may accumulate on the sorbent to levels in the rangeof from about 3,000 to 70,000 ppm of total metals, preferably 10,000 to30,000 ppm, of which 5 to 100%, preferably 20 to 80% is vanadium.Additives may be introduced into the contactor, the combustor or otherupgrading system components to passivate the non-selective catalyticactivity of these levels of heavy metals on the sorbent. Preferredadditives for this purpose include those disclosed in copending U.S.patent application Ser. No. 263,395, referenced above. Other metaladditives specifically for reducing the adverse effects of vanadium onsorbent may also be employed as described in PCT InternationalApplication Ser. No. PCT/US81/00357 filed in the U.S. Receiving Officeon Mar. 19, 1981, in the names of Ashland Oil, Inc., et al., andentitled "Immobilization of Vanadia Deposited on Sorbent MaterialsDuring Treatment of Carbo-Metallic Oils", the entire disclosure of saidPCT international application being incorporated herein by reference.

Referring to FIG. 2, a MRS feed is introduced into the bottom of aprogressive flow contactor 4 along with a suspension of hot sorbentprepared in accordance with this invention. Steam, naphtha, water, fluegas and/or some other diluent is preferably introduced into thecontactor along with the feed. These diluents may be from a fresh sourceor may be recycled from a process stream in the refinery. Where recyclediluent streams are used, they may contain hydrogen sulfide and othersulfur compounds which may help passivate adverse catalytic activity byheavy metals accumulating on the sorbent. It is to be understood thatwater diluents may be introduced either as a liquid or as steam.Diluents are added primarily as a source of vapor for dispersing thefeed and/or accelerating the feed and sorbent to achieve the vaporvelocity and residence time desired. Diluents other than water need notbe added but where used, the total amount of diluent specified includesthe amount of water used. Increases in diluent will further increase thevapor velocity and further lower the feed partial pressure in the riser.

As the feed travels up the contactor 4, it forms basically four productsknown in the industry as dry gas, wet gas, naphtha, and RCC or FCC feed.The naphtha is of relatively low octane and either may be sent to areforming operation to increase its octane number or combined with theRCC or FCC feed to form a blend which may be cracked and reformed in asubsequent conversion unit. At the upper end of the contactor, thesorbent particles are ballistically separated from product vapors. Thesorbent which then contains the coke formed in the contactor is sent tothe combustor 11 to burn off the coke and the separated product vaporsare sent to a fractionator (not shown) for further separation andtreatment to provide the four basic products indicated. The preferredconditions for contacting feed and sorbent in the contactor aresummarized in Table 5, in which the abbreviations used have thefollowing meanings: "Temp." for temperature, "Dil." for diluent, "pp"for partial pressure, "wgt" for weight, "V" for vapor, "Res." forresidence, "S/O" for sorbent to oil ratios, " sorb." for sorbent, "bbl"for barrel, "MAT" for micro-activity by the MAT test using a standardDavison feedstock, "Vel." for velocity, "cge" for charge, "d" fordensity and "Reg." for regenerated.

                  TABLE 5                                                         ______________________________________                                        Sorbent Riser Conditions                                                                     Board                                                                         Operating     Preferred                                        Parameter      Range         Range                                            ______________________________________                                        Feed Temp.     400-800° F.                                                                          400-650° F.                               Steam Temp.    20-500° F.                                                                           300-400° F.                               Reg. Sorbent Temp.                                                                           900-1600° F.                                                                         1150-1500° F.                             Riser Exit Temp.                                                                             800=1400° F.                                                                         900-1100° F.                              Pressure       0-100 psia    10-50 psia                                       Water/Feed     0.01-0.30     0.04-0.15                                        Dil. pp/Feed pp                                                                              0.25-3.0      1.0-2.5                                          Dil. wgt/Feed wgt                                                                            ≦0.4   0.1-0.3                                          V. Res. Time   0.1-5 sec.    0.5-3 sec.                                       S/O, wgt       3-18          5-12                                             Lbs. Sorb./bbl Feed                                                                          0.1-4.0       0.2-2.0                                          Inlet Sorb. MAT                                                                              <25 vol. %    <20 vol. %                                       Inlet Sorb. Rel. Act.                                                                        <1%           <0.5%                                            Outlet Sorb. MAT                                                                             <20 vol. %    <10 vol. %                                       Outlet Sorb. Rel. Act.                                                                       <0.5%         <0.1%                                            V. Vel.        ≧25 ft./sec.                                                                         ≧30 ft./sec.                              V. Vel./Sorb. Vel.                                                                           ≧1.0   1.2-2.0                                          Dil. Cge. Vel. 5-90 ft./sec. 10-50 ft./sec.                                   Oil Cge. Vel.  1-50 ft./sec. 5-50 ft./sec.                                    Inlet Sorb. d  1-9 lbs./ft..sup.3                                                                          2-6 lbs./ft..sup.3                               Outlet Sorb. d 1- 6 lbs./ft..sup.3                                                                         1-3 lbs./ft..sup.3                               ______________________________________                                         Sorbent particle circulation and operating parameters are brought up to     process conditions by methods well-known to those skilled in the art. The     equilibrium sorbent at a temperature of 1,150°-1,500° F.     contacts the oil feed at riser wye 17. The feed can contain steam and/or     flue gas injected at point 2 or water and/or naphtha injected at point 3     to aid in feed vaporization, sorbent fluidization and controlling contact     time in contactor riser tube 4. The feed is passed through contactor 4 in     contact with the hot fluidized inert sorbent to effect removal of metal     contaminants and coke precursors by thermal degradation of high boiling     and/or non-vaporizable materials in the feed. The liquid feed product     provided by thermal degradation of these materials also has a lower     viscosity so that this process may be referred to as a viscosity breaking     or "visbreaking" process.

The sorbent and vaporous hydrocarbons travel up contactor 4 at a contacttime of 0.1-5 seconds, preferably 0.1-3 seconds and more preferably0.5-2 seconds. At the upper end of contactor 4, the product vapors areballistically separated from the sorbent particles in vessel 5. Theriser tube of contactor 4 is of the vented type having an open upper end40 surrounded by a cup-like member 42 which preferably stops just belowthe upper end 40 of the riser so that the lip of the cup is slightlyupstream of the open riser tube as shown in FIG. 2. A product vapor line44 communicates with the interior of the cup so as to discharge productvapors entering the cup from the vapor space of vessel 5. The cup formsan annulus 46 around and concentric to the upper end of the riser tube.The transverse cross-sectional area of annulus 46 is preferably in therange of 70% to 120%, more preferably about 100%, of the transversecross-sectional area of the riser tube. This structure causes productvapors to undergo almost a complete reversal in their direction of flowupon being discharged from the riser tube and before leaving the vaporspaced of vessel 5. The product vapors then make a further turn orchange in direction of about 90° as they enter product line 44. Theproduct vapors then enter a cyclone separator 7 having an overhead vaporconduit 8. The amount of particle carry-over with this flow reversalstructure may be reduced by a factor of about 5 or more relative tocarry-over with the basic vented riser arrangement described in U.S.Pat. Nos. 4,066,533 and 4,070,159. Due to this reduction in carry-over,cyclone separator 7 may comprise only a single cyclone stage instead ofhaving multiple stages as is usually required to prevent excessivecarry-over of catalyst fines into the overhead vapor line in priorvented riser applications.

The sorbent and vaporous hydrocarbons are separated at vented riseroutlet 40 at a final contact temperature of about 900°-1100° F. Thevaporous hydrocarbons are transferred to cyclone 7 via cup 42 andconduit 44 where any entrained sorbent fines are separated and thehydrocarbon vapors are sent to a fractionator (not shown) via vaportransfer line 8. The coked sorbent is transferred to coked sorbentstripper 10 for removal of entrained hydrocarbon vapors. One method thatmay be employed in stripping residual hydrocarbons from coked sorbent isdescribed in international patent application Ser. No. PCT/US81/00646filed on May 13, 1981, in the name of Ashland Oil, Inc., et al., andentitled "Stripping Hydrocarbons from Catalyst with Combustion Gases",the entire disclosure of said PCT application being incorporated hereinby reference.

Stripped sorbent is then transferred to combustor vessel 11 to form adense fluidized bed 12. An oxygen containing gas such as air is admittedto the bottom of dense bed 12 by conduit 14 to combust the coke tocarbon oxides. The regenerating gas may be any gas which can provideoxygen to convert carbon to carbon oxides. Air is highly suitable forthis purpose in view of its ready availability. The amount of airrequired per pound of coke for combustion depends upon the desiredcarbon dioxide to carbon monoxide ratio in the effluent gases and uponthe amount of other combustible materials present in the coke, such ashydrogen, sulfur, nitrogen and other elements capable of forming gaseousoxides at regeneration conditions. The combustor vessel illustrated inFIG. 2 is a simple one zone-dense bed type. This regenerator apparatusis not limited to this example but may comprise two or more zones instacked or side by side relation and with internal and/or externaltransfer lines from zone to zone. Such multistage regenerators aredescribed in more detail in Ashland's RCC applications as referencedherein.

The combustor is operated at temperatures in the range of about 900° to1,600° F., preferably 1,150° to 1,500° F., to achieve adequatecombustion while keeping sorbent temperatures below those at whichsignificant sorbent degradation can occur. In order to control thesetemperatures, it is necessary to control the rate of burning which, inturn, can be controlled at least in part by the relative amounts ofoxidizing gas and carbon introduced into the regeneration zone per unittime. With reference to FIG. 2, the rate of introducing carbon into thecombustor may be controlled by regulating the rate of flow of cokedsorbent through valve 26 in conduit 28, the rate of removal ofregenerated sorbent by regulating valve 30 in conduit 16, and the rateof introducing oxidizing gas by the speed of operation of blowers (notshown) supplying air to the conduit 14. These parameters may beregulated such that the ratio of carbon dioxide to carbon monoxide inthe effluent gases is equal to or less than about 4.0, preferably equalto or less than about 1.5. In addition, water, either as liquid orsteam, may be added to the combustor to help control temperatures and toinfluence the carbon dioxide to carbon monoxide ratio. A particularlypreferred mode of operating the combustor so as to control the CO₂ /COratio and regenerate sorbent in a manner which helps minimize theadverse effects of vanadia is disclosed in U.S. patent application Ser.No. 255,931 filed on Apr. 20, 1981, in the name of William P. Hettinger,Jr., et al., and entitled "Immobilization of Vanadia Deposited onSorbent Materials During Treatment of Carbo-Metallic Oils", the entiredisclosure of said application being incorporated herein by reference.

The regenerator combustion reaction is carried out so that the amount ofcarbon remaining on regenerated sorbent is less than about 0.50 percent,preferably less than about 0.25 percent and most prefereably less thanabout 0.10 percent on a substantially moisture-free weight basis. Theresidual carbon level is ascertained by conventional techniques whichinclude drying the sorbent at 1,100° F. for about four hours beforeactually measuring the carbon content so that the carbon level obtainedis on a moisture-free basis.

The flue gas resulting from sorbent regeneration is processed through amulti-stage cyclone 22 and exits from combustor vessel 11 via line 23.The regenerated sorbent is transferred to regenerated sorbent stripper15 to remove any entrained combustion gases and then transferred toriser wye 17 via line 16 to repeat the processing cycle.

At such time as the heavy metals level (Ni+V+Fe+Cu) on the sorbentbecomes intolerably high so that sorbent effectiveness declines undulyand/or catalytic cracking activity increases unduly, additional sorbentcan be added and deactivated sorbent withdrawn at addition-withdrawalpoint 18 into dense bed 12 of combustor 11 and/or at addition-withdrawalpoint 19 into regenerated sorbent standpipe 16. Addition-withdrawalpoints 18 and 19 can be utilized to add either untreated used catalystor used catalyst already treated so as to contain one or more alkalinemetal additives of the invention.

In the case of a catalytically active sorbent without additive, thealkaline metal additive may be introduced into the sorbent material asan aqueous or hydrocarbon solution or as a volatile compound during theprocessing cycle. The alkaline metal additive as an aqueous solution oras an organo-metallic compound in an aqueous and/or hydrocarbon solventis preferably added directly to the sorbent at points 18 and 19, and/ormixed with the feed at addition points 2 and 3 into feed line 1.Addition point 20 into contactor 4 and addition point 21 into sorbentbed 9 in vessel 5 may also be employed for this purpose. The addition ofthe alkaline metal additive is not limited to these locations, but canbe introduced at any point in the oil/sorbent processing cycle. Withreference to FIG. 2, this would include, but not be limited to, additionof the alkaline metal additive solution at the contactor wye 17,anywhere along the length of contactor 4, to the dilute phase in sorbentseparator vessel 5, to the strippers 10 and 15, to the combustor airinlet 14, to the combustor dilute phase, and/or anywhere alongregenerated sorbent standpipe 16.

COMBINED MRS AND RCC PROCESSES

The vanadia immobilizing sorbents and the metals-Conradson carbonremoval processes described in this specification are preferablyemployed in combination with processes and apparatuses forcarbo-metallic oil conversion (RCC) of the type described in co-pendingU.S. applications Ser. Nos. 94,091, 94,092, 94,216, 94,217 and 94,227,all filed Nov., 14, 1979; and Ser. Nos. 246,751, 246,782 and 246,791,all filed March 23, 1981; said applications being in the name of GeorgeD. Myers alone or jointly with Lloyd E. Busch and the entire disclosureof each of said applications being incorporated herein by reference. Forexample, the spent catalyst from one of these RCC processes may betreated and used according to the present invention and the liquidproduct of the upgrading process of the invention may be used as a feedfor the same or a different RCC process. Other RCC conversion and/orcatalyst regeneration processes that may be combined with the upgradingprocesses of the present invention include those described in U.S.applications Ser. No. 263,398, filed on May 13, 1981, and Ser. No.258,265 filed on Apr. 28, 1981 , both in the name of William P.Hettinger, Jr., and others, and in international application Ser. No.PCT/US81/00662 filed in the name of Ashland Oil, Inc., et al., on May13, 1981, the entire disclosures of said U.S. and internationalapplications being incorporated herein by reference.

The preferred feeds capable of being cracked by these RCC methods andapparatuses are comprised of 100% or less of 650° F.+ material of whichat least 5 wt%, preferably at least 10 wt%, is of high molecular weightand does not boil below about 1,025° F. The terms "high molecularweight" and/or "heavy" hydrocarbons refer to those hydrocarbon fractionshaving a normal boiling point of at least 1,025° F. and includenon-boiling hydrocarbons, i.e., those materials which may not boil underany conditions.

A carbon-metallic feed for purposes of RCC processing is one having aheavy metal content of at least about 4 ppm nickel equivalents (ppmtotal metals being converted to nickel equivalents by the formula: NiEq.=Ni+V/4.8+Fe/7.1+Cu/1.23) and a Conradson carbon residue valuegreater than about 1.0. The upgraded feeds for which RCC processing isparticularly useful will have a heavy metal content of at least about 5ppm nickel equivalents, a vanadium content of at least 2.0 ppm, and aConradson residue of at least about 2.0.

With respect to the tolerance levels of heavy metals on RCC crackingcatalysts, such metals may accumulate on these catalysts to levels inthe range of from about 3,000 to 70,000 ppm of total metals, preferably10,000 to 30,000 ppm and more preferably 15,000 to 20,000 ppm, of whichpreferably 5 to 100%, more preferably 20 to 80%, may be vanadium.

Referring now to FIG. 3 by way of example, there is shown a simplifiedblock flow arrangement of a combination process according to theinvention. In the arrangement of FIG. 3, a high boiling hydrocarbon feedsuch as a reduced crude is charged by conduit 102 to a MRS upgradingsystem 104 wherein the feed is contacted with the solid sorbentparticulate material of the invention to thermally visbreak the feed soas to reduce its Conradson carbon value and metal contaminants to alower, more acceptable level. During this thermal contacting operationwith solid sorbent particulates, a wet gas product is formed andrecovered by conduit 106, a C₅ + naphtha product is formed and recoveredby conduit 108 and a 430° F.+ product is formed and recovered by conduit110.

The C₅ plus product material in conduit 108 and the 430° F.+ productmaterial in conduit 110 are charged to an RCC (reduced crude cracking)unit 112 more fully discussed below with respect to FIGS. 4 and 5.During catalytic cracking conversion of the demetallized feed materialobtained from the MRS operation, a wet gas product stream is formed andrecovered by conduit 114, a main column overhead liquid is recovered byconduit 116, a 430° to 630° F. product fraction is recovered by conduit118 and a 630° F.+ material stream containing both product andunconverted feed is recovered by conduit 120. The RCC conversionproducts recovered by conduits 118 and 120 may be used in fuel oilblending operations not shown. On the other hand, the 630° F.+ productmay be further converted by catalytic cracking either in a separate FCCunit not shown or as recycle to the RCC unit 112. The 430° F. to 630° F.product in conduit 118 may also be further refined as desired to meetfuel demands as by cracking, hydrogenation, reforming and/or otherprocessing steps suitable for the product desired.

The RCC main column overhead liquid in conduit 116 is passed to a gasconcentration unit 122. The wet gases in conduits 106 and 114 areblended and also passed to unit 122 by conduct 107. The operation of thegas concentration unit provides a fuel gas stream withdrawn by conduit124, which material is then passed to an MEA absorber 126 before beingrecovered by conduit 128.

A gasoline product boiling in the range of C₅ hydrocarbons up to about400° to 430° F. is recovered from unit 122 by conduit 130 and passed toa gasoline treating unit 132 before being recovered by conduit 134. Ingasoline treating unit 132, it is contemplated treating the gasolinewith a caustic wash and an electrostatic precipitator to removeundesired impurities in a manner known in the industry.

A C₃ /C₄ product fraction is recovered from unit 122 by conduit 136 andpassed to a C₃ /C₄ treating unit 138. In unit 138, the C₃ /C₄ fractionis treated to remove sulfur impurities and then caustic washed. Theproduct of treating unit 138 is passed by conduit 140 to a hydrogenationunit 142 designed to particularly accomplish hydrogenation of diolefinsbefore being recovered by conduit 144.

Used equilibrium catalyst is withdrawn in a fully regenerated conditionfrom the regenerated standpipe of the RCC unit and transferred via line150 to a treatment unit 152 for impregnating the catalyst with analkaline metal additive to neutralize its residual catalytic activity.The catalyst is preferably soaked in an aqueous solution of the selectalkaline metal additive, dried and then transferred to the MRS unit 104as makeup sorbent via line 154.

Referring now to FIG. 4, by way of example there is shown onearrangement of apparatus for effecting thermal contact of the residualoil or reduced crude containing feed with the solid sorbent material toaccomplish metals removal and reduce the Conradson carbon producingcomponents of the feed by thermal visbreaking. The contact system ofFIG. 4 is referred to herein as the MRS system (metals removal system)and comprises a riser contact zone for selectively contacting the heavyresidual oil feed with hot solid sorbent particulates comprisingneutralized catalyst of little or no catalytic cracking activity. Thecontact temperature, space velocity conditions and hydrocarbon feedpartial pressure are selected to accomplish substantial metals removalin the absence of excessive thermal cracking for the production ofthermal naphtha and higher boiling range products.

In the specific arrangement of FIG. 4, atomizing water is added byconduit 101 to a reduced crude containing feed introduced by conduit 103to a riser contactor 105 above the bottom portion thereof. Steam inconduit 97 and/or admixed with water in conduit 96 obtained from themain column overhead drum is admixed with regenerated hot solid sorbentat a temperature in the range of 1300° to 1500° F. in the bottom portionof the riser in respective amounts and under conditions selected toadjust the temperature of the hot sorbent before it contacts the oilfeed charged to the contactor. This particular combination of diluentsadmixed with solids permits establishing a vertical velocity componentin the solids before they contact dispersed hydrocarbon feed materialfurther along the riser. The suspension of sorbent particulates andatomized feed of low partial pressure in the presence of steam diluentis passed through the riser contact zone at a temperature below about1025° F. and at a velocity providing a hydrocarbon residence time ofless than 5 seconds and preferably within the range of 0.5 to 4 seconds.The contactor 105 is provided with a plurality of vertically spacedapart feed inlet means to accomplish desired changes in hydrocarbonresidence time (not shown).

The suspension passed through contactor 105 is discharged from the topor open end of the riser and vaporous hydrocarbons from thermalvisbreaking and gasiform diluent material are ballistically separatedfrom sorbent particulates in the manner previously described withreference to FIG. 2 and are caused to flow through a plurality ofparallel arranged cyclone separators 111 and 113 positioned about theupper open end of the riser contact zone. Hydrocarbon vapors separatedfrom entrained solid by the cyclone separators are collected in a plenumchamber 115 before withdrawal or recovery by conduit 117 at atemperature of about 970° F. The vaporous material in conduit 117 isquenched in one specific embodiment to a temperature of about 680° F. byadmixture with a portion of the bottoms product from the main column ofthe MRS fractionator (not shown) which is introduced into conduit 117through conduit 119.

Solid particulate material comprising the select sorbent of thisinvention and accumulated metal deposits and carbonaceous material fromthermal degradation of the feed are collected in a bottom portion ofsorbent separator vessel 121 which includes a stripping section 123 towhich stripping gas is charged by conduit 125 at a temperature of atleast about 400° F. Higher stripping temperatures up to about 1050° F.are also contemplated.

Stripped solid sorbent material is passed by standpipe 127, which isprovided with a flow control valve, to a fluid bed of sorbentparticulates in a bottom regeneration zone 129 of a combustor 152.Regeneration gas or combustion supporting gas, such as oxygen modifiedgas or air, is charged to this lower regeneration zone by a conduit 131and through a plenum distribution chamber 133 supporting a plurality ofradiating gas distributor pipes 135. Regeneration of the sorbentparticles by burning off the deposited carbonaceous material isaccomplished at a temperature up to about 1500° F. and preferably in anoxygen lean or restricted oxygen containing atmosphere promoting theformation of a carbon monoxide rich regeneration flue gas. Combustionproduct gases and sorbent pass from a dense fluild bed 137 ofparticulates through a restricted passageway 139 as a suspension ofparticulates in flue gases to an upper enlarged portion of the combustorvessel where a ballistic separation is made between solid particulatesand combustion flue gases by the combination of a ballistic separator 49of the type described in reference to FIG. 2 and cyclone separators43,43 positioned about the open upper end of passageway 139. Theseparated particulate is collected as a fluid bed of material 141 in anannular zone about restricted passageway 139. Flue gases separated fromthe solids pass through the annular cup of ballistic separator 49 andinto the cyclones 43, 43 for removal of entrained fines. The CO richflue gases then pass to plenum chamber 45 for discharge through conduit47. Regenerated sorbent at an elevated temperature within the range of1000° to 1500° F. is passed by standpipe 149 to the bottom portion ofriser 105 for use as herein described. A portion of the hot regeneratedsorbent may be withdrawn by a conduit 151 for passage to a heatexchanger 153 wherein the sorbent may be cooled and steam generated byindirect heat exchange with boiler feed water introduced by conduit 155.The steam is recovered by conduit 157. The thus partially cooled solidparticulate may be withdrawn by conduit 159 for passage to a bottomportion of the fluid bed of particulate in a bottom portion ofregeneration zone 129 for temperature control of the particulate solidsbeing regenerated.

Alternately, a slip stream of cooled sorbent may be transferred byconduit 161 to a soak vessel 163 for impregnation with a solution of theselect alkaline metal additive introduced through line 165. Afterimpregnation, the additive solution is drawn off via line 167 and thesorbent dried with hot combustion air introduced by line 169 and drawnoff by vent and/or surge line 171. Dry sorbent freshly treated withalkaline metal is then returned to the combustor via a standpipe 173 andan air lift line 175.

Referring now to FIG. 5 there is shown an arrangement of vessels withinterconnecting conduits comprising an RCC riser reactor-regeneratorsystem relied upon to process the demetallized feed materials obtainedfrom the MRS upgrading unit. As mentioned above, the residual metals inthe RCC feed on the basis of nickel plus vanadium (Ni+V) charged to theRCC riser reactor will be considerably reduced by the MRS solidscontacting unit and usually will be less than about 100 ppm, preferablyless than 50 ppm and more preferably less than 25 ppm Ni+V.

The operation of the RCC system is similar in many respects to a fluidcatalyst cracking (FCC) system but is more critically and selectivelycontrolled with respect to catalyst regeneration and catalyst use inorder to process carbo-metallic oils, such as reduced crudes, which havebeen partially demetallized and decoked as herein provided. Operation ofthe RCC unit is carefully controlled with respect to heat balance andfeed conversion selectivity since relatively large amounts ofcarbonaceous deposits (coke) are laid down on the catalyst depending onthe level of contaminates in the feed charge and the amount of metaldeposits accumulated on the circulated catalyst. In addition,overheating and degradation of the catalyst and non-selective and/orover cracking of the feed are potentially eminent so that productselectivity may be undesirable altered when the RCC operation ispermitted to vary from a relatively restricted range of operatingconditions into less than an efficient and economic operation. Thus itis desirable to utilize a two stage regeneration operation in the RCCsystem so as to restrain the amount of heat transferred to the RCC risercracking zone by the catalyst and yet reduce residual coke onregenerated catalyst to a desired low level, preferably less than 0.05wt% and more preferably less than 0.01 wt%.

A particularly important feature is the identification of operatingconditions in the hydrocarbon conversion riser reactor which permitprocessing upgraded heavy crudes under conditions particularlyincreasing the yield of desired liquid fuel products such as gasolineand gasiform materials readily converted to useful liquid fuel productsby auxiliary processes such as alkylation, isomerization, polymerizationor a combination thereof. Both the MRS and RCC riser operations arerelatively high velocity operations as herein described. In the RCCsystem, the temperature of the conversion product vapors as measurednear the riser outlet is preferably restricted to within a range of 950°F. to about 1050° F. and more preferably restricted to less than about1000° F. when processing demetallized feed from the MRS unit. The MRSfeed is charged to the RCC riser at a temperature of about 400° F. Theuse of water injection and steam dispersion is relied upon insubstantial measure for temperature control and feed atomization and forachieving a high velocity suspension of catalyst particles andhydrocarbon feed. This velocity and the riser dimensions are such thatthe hydrocarbons are in contact with catalyst particles in the riserconversion zone for less than about 4 seconds, preferably less thanabout 3 seconds and more preferably less than about 2 seconds. Thus itis important to achieve rapid dispersion of feed, rapid contact ofdispersed feed with catalyst particles and rapid separation ofhydrocarbon vapors from catalyst substantially immediately upondischarge of the suspension from the riser conversion zone. In order toaccomplish these objectives, the riser reactor is designed to achievehigh velocity (preferably 180 ft./second or more) mixing of hot catalystparticles with fluidizing gaseous material and with charged feedmaterial in a lower, restricted diameter portion of the riser reactor.In an upper, larger diameter portion, the velocity of the suspension isdescreased to a discharge velocity of about 85 to 95 feet per second orless adjacent the upper discharge end of the riser. In the embodiment ofFIG. 5, this alteration of the suspension velocity is accomplished bygradually increasing the diameter of the riser in an upper transitionsection thereof. In a particularly preferred arrangement, the initiallyformed suspension is at a velocity of about 180 feet per second and thevelocity of the suspension discharged from the open upper end of theriser reactor is about 86 feet per second after passing through thegradually expanding transition section. Preferably, the dischargetemperature is maintained below about 1025° F. but above about 975° F.in this specific embodiment. In the specific arrangement of FIG. 5, theriser reactor is about 198 ft. tall, of which the top 25 feet thereof isabout 63 inches in diameter and connected by a 7 foot transition sectionto a 40 inch diameter riser tube comprising the lower section of theriser reactor.

A mixture of the demetallized 430° F. plus feed obtained from the MRSunit, the MRS naphtha and liquid water obtained from the main columnoverhead drum is introduced into a bottom or upper portion of the riserreactor 50 as desired using charge lines 51 or 53, respectively. Steammay be charged via line 55 to the bottom of the riser, either separatelyor in admixture with demetallized feed charged via line 57, to assistwith atomization and/or dispersion of the heavy oil feed which contactsregenerated hot catalyst particles charged to the bottom portion of theriser by conduit 52. It is thus contemplated forming a high velocitysuspension of steam and catalyst, with or without charged naphtha, in abottom portion of the riser prior to charging the demetallized,thermally converted oil feed which may be mixed with additionalquantities of water and/or steam. It is further comtemplated chargingthe oil feed admixed with atomizing diluent such as steam and/or naphthathrough a plurality of oil feed atomizing devices (not shown) to achievemore intimate contact of the high boiling feed with hot regeneratedcatalyst particles. These materials form a suitable upflowing suspensionin which the feed undergoes conversion to the liquid fuel productsdesired within the residence time and at the conversion conditionsprovided. The concentration of catalyst particles in the upflowingsuspension is preferably within the range of about 2 to 10 lbs. percubic ft.

Upon transversing riser 50 and being discharged from its top open end,product vapors are abruptly separated from catalyst by ballisticseparator 59 which is of the type described above in reference to FIG.2. After reversing direction and passing through separator 59, thegasiform product material comprising hydrocarbon vapors with someentrained catalyst particles passes through a plurality of parallelarranged cyclone separators 54, generally comprising one stage ofcyclone separation, positioned about the upper end of the riser.However, two or more cyclone separators or stages in series may also beemployed. Separated product vapors are collected in plenum chamber 56and withdrawn therefrom by conduit 58 communicating with an RCC productrecovery section which includes a main fractionator column (not shown).Catalyst particles separated from the vaporous products of hydrocarbonconversion are collected and passed downwardly through a stripping zone60 in countercurrent relation to a stripping gas such as steamintroduced by conduit 62. The temperature of the RCC stripping operationis generally maintained within the range of 900° F. to 1050° F. Thestripped catalyst is passed by a standpipe conduit 64 to a bed ofcatalyst 66 in a first stage of catalyst regeneration affected in theupper portion of a two stage catalyst regenerator 68.

The upper stage of regenerator 68 is of larger diameter than the lowerstage and the two stages are separated from one another by a centrallylocated air distributor having a plenum chamber 70 supported by anannular baffle 72 provided with flue gas flow through passageways 74,74.A plurality of hollow radiating arms 76 are provided on plenum 70 fordistributing fresh regeneration gas to a lower portion of catalyst bed66. In a preferred embodiment, regeneration air amounting to about 75%of that required to achieve removal of carbonaceous material to thedesired level by burning is introduced to plenum chamber 70 by conduit78. The regeneration air or other oxygen containing gas is preferablypreheated by means not shown to a desired elevated temperature of about300° F. The catalyst introduced to bed 66 by conduit 64 comprisescarbonaceous deposits from the riser hydrocarbon conversion process andthese deposits contain significant amounts of hydrogen which is oxidizedto water at regeneration conditions.

In the first stage of catalyst regeneration, the regenerationtemperature is kept to a low value, preferably in the range of about1200° F. up to about 1300° F. under combustion supporting conditionssufficient to effect at least partial removal of the carbonaceousdeposits and produce a CO rich flue gas. The CO rich flue gases,comprising CO₂, sulfur oxides, nitrogen and water vapor, pass from thedilute phase of the first stage through a combination of parallelarranged cyclone separators which may comprise two or more stages inseries and are represented by cyclones 80 in which entrained catalystparticles are separated from flue gas and returned by the cyclone diplegto catalyst bed 66. The CO rich flue gas separated from catalystparticles by the cyclones is passed to a plenum chamber 82 forwithdrawal therefrom by a conduit 84.

The partially regenerated catalyst of bed 66 is passed to a second stageof catalyst regeneration below the first stage through an externaldowncomer 86 provided with a catalyst cooler 88 wherein high pressure(450 psig) steam may be generated. The catalyst, partially cooled incooler 88 and only partially regenerated, is then passed by conduit 90to a fluid bed of catalyst 92 in the lower portion of regenerationvessel 68. Regeneration of the catalyst is completed in bed 92 whichcomprises the second stage of catalyst regeneration. A standpipe 94 isalso provided as a second external standpipe for transfer of catalystfrom bed 66 to bed 92 without water cooling. However standpipe 86 is theprimary route for catalyst transfer from bed 66 to bed 92. The purposeof these transfer standpipes is to transfer partially regeneratedcatalyst to and maintain temperature control in the lower regenerationstage. In this second stage, residual carbon on the catalyst ispreferably burned at a temperature within the range of about 1325° up toabout 1500° F. Regeneration of catalyst in bed 92 is therefore affectedat a temperature generally higher than that of bed 66. Bed 92 is morepreferably maintained at a temperature within the range of about 1350°to about 1400° F. and in the presence of sufficient excess oxygen toachieve very low levels of residual coke on the partially regeneratedcatalyst received from bed 66.

In one preferred embodiment, the amount of air or oxygen modified gasemployed in the second stage of catalyst regeneration is only about 25%of that required to burn all of the coke on the catalyst entering theregenerator from line 64. This second stage regeneration gas isintroduced beneath a distribution grid 93 by a conduit 91. All of theflue gas from the second stage of regeneration passes through openings74 in baffle member 72 separating the upper stage from the lower stage.Thus the hotter flue gases of the second stage comprising CO, CO₂ andexcess or unreacted oxygen pass into the bottom portion of bed 66 andthereby contribute heat to catalyst bed 66 and help initiate thecombustion of carbonaceous deposits on incoming catalyst from conduit64. Regenerated catalyst of relatively low residual coke, preferablybelow about 0.10 and more preferably below 0.05 weight percent, and at atemperature within the range of 1300° to 1500° F. is withdrawn fromcatalyst bed 92 for passage by standpipe 52 to a lower bottom portion ofriser 50 and reuse as hereinbefore discussed.

Although used FCC catalysts may be employed as the MRS sorbent and maybe preferred where they contain less heavy metal contaminants than acomparable RCC catalyst, the present invention contemplates integratedMRS and RCC systems wherein the MRS sorbent comprises an RCC catalystwithdrawn from the integrated RCC system. One embodiment of such an RCCsystem is illustrated in FIG. 5 wherein a side stream of cooled,partially regenerated catalyst is withdrawn and transferred by a conduit61 to a treatment vessel 63 for impregnation with a solution of theselect alkaline metal additive introduced into the vessel by line 65.After the catalyst has been immersed in the additive solution for thedesired period of time, the solution is drawn off through line 67 andthe resulting sorbent dried with hot combustion air introduced throughline 69 and drawn off or vented through a vent or surge line 71. Drysorbent containing an amount of alkaline metal sufficient to neutralizesubstantially all of the catalytic sites is then transferred to the MRSsystem of FIG. 4 via a sorbent transfer line or standpipe 73. This freshmake-up sorbent may be added directly to either the upper or lower stageof combustor 152 or to regenerated sorbent standpipe 149. Alternately,if neutralization of the catalytic sites has not been substantiallycompleted in RCC treatment vessel 63, the partially neutralized catalystmay be transferred to vessel 163 in the MRS system and furtherimpregnated with the alkaline metal solution, MRS vessel 163 comprisingpart of the auxiliary treatment system previously described. Theinvention further contemplates introducing untreated or partiallytreated catalyst directly into combustor 152 and/or standpipe 149 andintroducing the alkaline metal solution with the oil feed of line 103and/or the water of lines 96, 97 and/or 101 in lieu of or in addition totreating a slip stream of untreated and/or partially treated sorbent inauxiliary vessel 163.

The alkaline metal addition vessels of FIGS. 4 and 5 include valving(not shown) in the associated conduits for isolating a batch ofcatalyst, treating it with an alkaline metal solution at near ambientconditions, and draining off the treatment solution before the catalystis transferred as sorbent to the MRS upgrading system. In this manner,undesirable components in the treating solution which may be detrimentaleither to the sorbent or to the upgrading process, such as exchangedrare earth metals or other agents capable of reactivating catalyticsites, are removed before the sorbent is added to the upgrading system.However, where the treatment solution does not contain components whichmight be detrimental to the sorbent or the upgrading process, thesolution may be added to a moving side stream wherein untreated orpartially treated catalyst is continuously drawn off from the RCCregenerator and/or the MRS combustor, treated and then continuouslytransferred as sorbent to the MRS system. In this alternative, the sidestream vessel 63 of FIG. 5 and the slip stream vessel 163 of FIG. 4would serve as surge vessels for contacting a fluidized volume offlowing catalyst with the treating solution as untreated and/orpartially treated catalyst continuously passes through these vessels.

When treating a continuous stream of hot catalyst, the liquid componentof the treating solution may vaporize upon contacting the hot catalystso as to deposit the metal additive on catalyst particles in thecontinuously moving stream. One advantage of this arrangement is thatthe heat of vaporization of the liquid can be used to help control thetemperature of the highly exothermic combustion reaction in thecombustor. One disadvantage of contacting very hot catalyst is that aportion of the treating solution may be solidified before the catalystparticles can be impregnated with the alkaline metal additive. Thecatalyst is therefore preferably cooled by RCC cooler 88 and/or MRScooler 153 to a temperature permitting uniform and effectiveimpregnation without premature solidification and/or non-uniformaccumulations of the additive materials. The additive materials thendecompose in the combustor or regenerated sorbent standpipe so as toneutralize the active acid sites of the catalyst.

The vaporous hydrocarbon products and diluent materials of the RCCreactor are withdrawn by conduit 58 and passed to the main column of afractionator unit (not shown) for separation by fractionation. Theoutput of the fractionator unit is usually comprised of at least sevenstreams, namely, dry gas, wet gas, gasoline boiling range productmaterial, sour water, light cycle oil (LCO), heavy cycle oil (HCO) andslurry or clarified oil. Any one or more of these output streams may berecycled to the RCC riser reactor as diluent or lift gas and/or as afurther feed component undergoing cracking and/or reforming at theconversion conditions employed. The invention further contemplatessubstituting an FCC riser reactor and regeneration system for the RCCriser reactor and regeneration system shown in FIG. 5. The output andrecycle streams of such FCC apparatuses and processes are the same as orsimilar to those of the RCC apparatuses and processes described herein.Another important aspect of the combination operation of this inventionabove discussed is concerned with the severity of the thermal upgradingrelied upon in the MRS operation to provide a suitable feed forefficient conversion with a crystalline zeolite containing catalyst inthe downstream RCC operation. In other words, the operating techniquesof this combination of operations are concerned with affecting thecatalytic conversion of a thermally prepared feed material which maystill comprise a relatively high level of metal contaminants with azeolite cracking catalyst. The crystalline zeolite is in admixture witha sorbent matrix material having a relatively large pore volume and poredimensions and this catalyst is capable of higher heavy metal loadingsthan previously thought useable in such cracking operations. In thisoperating combination, tailoring of the operating conditions employed inthe MRS feed preparation operation is made to provide a thermallyprocessed high boiling feed for a downstream catalytic conversion step.This interrelated tailoring of operating conditions for feed preparationand utilization is adjusted dependent upon the initial composition,initial boiling point and initial level of metal and Conradson carboncontaminants of the high boiling feedstock to be upgraded, and upon thecomposition, boiling point and level of metal and Conradson carboncontaminates in the upgraded feed to be transferred from the MRS unit tothe downstream conversion unit, which may be either an FCC unit or anRCC unit depending on these interrelated parameters. In this connection,whatever high boiling portion of a crude oil is to be upgraded as hereinprovided, thermal preparation of the feed may be accomplished underconditions permitting up to 50 ppm heavy metals and up to 8 wt%Conradson carbon to remain in a feed to be charged to an RCC catalyticcracking operation. It is also desirable within this processingarrangement to limit the production of thermally produced naphtha eventhough such thermal naphtha can be and is intended to be upgraded in thezeolite cracking operation along with the thermally demetallized higherboiling feed.

It is thus evident from the discussion above presented that thecombination operation of this invention is economically attractive sinceit permits the use of solid particulate of fluidizable particle size forboth catalytic cracking and non-catalytic upgrading under conditions ofhigh metals loading thereby reducing solids inventory replacement. Ofparticular interest is the finding that such solid particulate can beused to advantage for economically preparing and processing high boilingresidual portions of crude oils comprising Ni+V metal contaminants inexcess of 100 ppm and Conradson carbon values in excess of 10 wt% toprovide more desirable liquid fuel and lower boiling products.

EXAMPLE OF SORBENT PREPARATION

A zeolite containing catalytically active catalyst of the typepreviously described was charged to an RCC unit processing 200 barrelsper day of reduced crude containing about 100 ppm heavy metals, whichequated to about 45 nickel equivalents based on the formula describedelsewhere in this specification. After processing this feed continuouslyfor three weeks, the catalyst had dropped in activity, as measured bythe ASTM D3907-80 method, from a MAT relative activity of 100% down to aMAT relative activity of 15%. The metal content of this used, lessactive catalyst was about 4,000 ppm Ni, 3,100 ppm Fe, 9,900 ppm V, whichequates to a Ni equivalents of about 6,400 ppm metals, and the catalysthad a MAT relative activity of about 15%.

The used catalyst was then withdrawn from the RCC unit and impregnatedwith an alkaline metal salt as follows: To 1,244 grams of the used RCCcatalyst was added 2,000 ml. of deionized water. To this slurry wasadded 25 grams of sodium carbonate (about 0.6 wt% sodium based on weightof catalyst) and the mixture was heated to 100° F. with stirring. Afterapproximately 15 minutes, the excess solution was decanted and thecatalyst was dried at 250° F. for two hours and calcined in air at 1000°F. for two hours. The resulting sorbent was analyzed by atomicadsorption (AA) and found to contain 1.05 wt% sodium. The sorbent wasalso tested by the ASTM D3907-80 MAT activity test procedure and foundto have a MAT activity of 25 volume percent conversion, which equates toa MAT relative activity of 0.35% where the standard material has a MATrelative activity of 100%.

EXAMPLE OF FEED UPGRADING

Approximately 1168 grams of the select sorbent prepared as described inthe foregoing example was placed in a product distribution unit (PDU)for evaluation as a sorbent to convert a high Conradson carbon-metalscontaining reduced crude to a low Conradson carbon-metals containing FCCor RCC feed. The PDU is a fixed bed test unit containing a fluidizablematerial in which a relatively large quantity of reduced crude feedstockcan be processed to obtain sufficient liquid product to determine itsphysical properties. Over the 1168 grams of select sorbent was passed119 grams of a reduced crude containing 23 ppm Ni, 115 ppm V, 11 ppm Feand 23 ppm Na and having a Conradson carbon value of 7.28. The operatingparameters of the contact zone were a temperature of 1000° F., a 15second contact time between feed and sorbent and a sorbent to oil (S/O)ratio of 10. The resulting liquid product contained Ni, V, Fe and Na inamounts of less than 1 ppm each, and had a Conradson carbon value of1.3. This represents over 99% removal of the metals and over 80%reduction in Conradson carbon content. The liquid product prepared inthis test sequence is an excellent FCC feedstock and would be a premiumRCC feedstock.

COMPARISON TESTING OF SORBENTS

To compare MRS operations using the select sorbent of the invention withthose using prior art clay sorbents, a series of particulate "clumping"tests were performed. A clumping test is carried out as follows: Avanadia containing sorbent sample is placed in an individual ceramiccrucible, dried and calcined at 1400° F. in air for two hours using amuffle furnace. At the end of the calcining period, the crucible iswithdrawn from the muffle furnace and cooled to room temperature. Thesurface texture and flow characteristics of the sorbent particles arethen observed. Surface vanadia, while liquid at the calciningtemperature (1400° F.), will flow across the sorbent surface and causesorbent particles to coalesce when cooled down below its solidificationpoint. The degree of coalescence is a visual and mechanical estimationof particle fusion, namely, "flowing"--no change in flow characteristicsbetween virgin sorbent and used sorbent; "soft"--substantially all ofused sorbent is free flowing with a small amount of clumps easilycrushed to free flowing sorbent; "intermediate"--free flowing usedsorbent contains both free flowing particles and fused masses inapproximately a 1:1 ratio; and "hard"--substantially all of the usedsorbent particles are fused into a hard mass with very few free flowingparticles.

A kaolin clay, such as described in U.S. Pat. No. 4,263,128 and spraydried to yield microspherical particles in the 20 to 150 micron sizerange, had vanadia deposited upon it in varying concentrations by usingthe clay as a sorbent in the treatment of a reduced crude to lower itsvanadium and Conradson carbon values. Two process runs were made witheach process period extending for approximately 30 days. Samples of theclay sorbent were taken at varying vanadium levels during the two runs.With reference to FIG. 6, the solid dots and solid triangles,respectively, represent the measured levels of vanadium on the claysorbent samples from these two processing periods. After being withdrawnfrom the process, each sample was subjected to the clumping testdescribed above. A clay sample free of vanadia was also tested as acontrol. The surface texture and flow characteristics of the calcinedsamples were noted and the results of these clumping tests are reportedin Table 6A.

                  TABLE 6A                                                        ______________________________________                                        V.sub.2 O.sub.5                                                                             Surface      Flow                                               Concentration - ppm                                                                         Texture      Characteristics                                    ______________________________________                                        0             Free         Free flowing                                        1,000-10,000 Surface Clumped                                                                            Broke crust for                                                               free flowing                                       10,000-20,000 Surface Clumped                                                                            Total clumping                                                                no flow                                            ______________________________________                                    

A catalytically active cracking catalyst was prepared by methods similarto those described above to yield microspherical catalyst particles inthe 20 to 150 micron size range. This catalyst was utilized in an RCCprocess to lower its catalytic activity and then neutralized with analkaline metal additive to yield a sorbent according to the invention.Vanadia was deposited upon samples of this sorbent in varyingconcentrations to simulate samples from an upgrading process and each ofthese samples was subjected to the clumping test described above. Asample of this sorbent free of vanadia was also tested as a control. Thesurface texture and flow characteristics of the calcined samples werenoted and the results of these clumping tests are reported in Table 6.

                  TABLE 6B                                                        ______________________________________                                        V.sub.2 O.sub.5                                                                             Surface      Flow                                               Concentration - ppm                                                                         Texture      Characteristics                                    ______________________________________                                        0             Free         Free flowing                                       15,000        Free         Free flowing                                       20,000        Free         Free flowing                                       30,000        Surface clumped                                                                            Broke crust for                                                               free flowing                                       40,000        Surface clumped                                                                            Partial flowing                                    50,000        Surface clumped                                                                            Total clumping                                                                No flowing                                         ______________________________________                                    

A comparison between the results obtained with the select sorbent ofthis invention and those obtained with the conventional clay sorbent isshown in FIG. 6. As also shown in Table 6A, the clay free of vanadiadoes not form any crust or clumps or fused particles at the temperaturesencountered in the regeneration zone of the upgrading process of theinvention. At vanadia concentrations of 1,000-10,000 ppm, clumping wasobserved but the vanadia crusts binding paarticles together could bereadily broken into free flowing, crusty particles. The conventionalclay sorbent particles began to show significant coalescence propertiesat vanadium levels of about 10,000 ppm and above, and by 20,000 ppm hadcoalesced into a hard mass evidencing a complete loss of fluidizationproperties.

Manifestation of the foregoing phenomena is further demonstrated by thefinding that when coalesced clay particles of high vanadium content(10,000 ppm or more) are cooled down from a temperature above thesolidification temperature of vanadium pentoxide in an operating MRSunit in order to permit entrance into the unit for cleaning out pluggeddiplegs and other repairs, a solid mass of sorbent which must beforcibly removed has been observed. This phenomena makes turn-aroundlengthy and complex for the operating unit since this material must bechipped out.

In comparison, significant particle coalescence of the select sorbent ofthe invention does not begin until a vanadium level of about 25,000 to30,000 ppm is reached, as opposed to significant coalescence of theconventional sorbent at about 10,000 ppm vanadium. Whereas theconventional clay sorbent is totally clumped and not free flowing at20,000 ppm V, the sorbent of the invention is still free flowing at30,000 ppm and does not experience total clumping until about 50,000 ppmV.

INDUSTRIAL APPLICABILITY

The invention is useful in the treatment of both FCC and RCC feedstocksas described above. It is particularly useful in the treatment of highboiling, carbo-metallic feedstocks of extremely high metals and/orConradson carbon values to provide products of lower metals and/orConradson carbon values suitable for use as feeds for FCC and/or RCCunits. Examples of these oils are reduced crudes and other crude oils orcrude oil fractions containing metals and/or residua as above defined.

Although the upgrading process is preferably conducted in a riserreactor of the vented type, other types of risers and other types ofreactors with either upward or downward flow may be employed. Forexample, the upgrading operation may be conducted with a moving bed ofsorbent which moves in countercurrent relation to liquid (unvaporized)feedstock under suitable contact conditions of pressure, temperature andweight hourly space velocity. The process conditions, sorbent and feedflows and schematic flow arrangement of such a moving bed operation aredescribed in the literature, such as for example, in the articleentitled, "T. C. Reforming", Pet. Engr., April 1954; and in the articleentitled "Hyperforming", Pet. Engr., April 1954; the entire disclosuresof said articles being incorporated herein by reference.

The economic advantages of the invention include the utilization ofcheap materials, such as deactivated, spent or equilibrium FCC or RCCcatalysts valued at $100 to $300 per ton, instead of freshly preparedclay sorbents valued at $800 to $1,000 per ton.

In the upgrading of various feeds, the rate of vanadium buildup on thesorbent and the equilibrium or steady state level of vanadium on thesorbent are functions of the vanadium content of the feed and especiallythe sorbent replacement rate at equilibrium conditions. Table 7 presentsa typical case for a 40,000 bbl/day MRS unit in which the vanadiumcontent of the feed in varied from 1 ppm, representative of upgrading anFCC feedstock of VGO and 5 to 20 percent of a heavy hydrocarbonfraction, up to a range of 25 to 400 ppm, representative of upgrading anRCC feedstock comprised of 70 to 100 percent of a reduced crude. Inorder to maintain various levels of vanadium on the sorbent at anequilibrium state achieved after long term operation (50 to 150 days),this sorbent replacement rate may be varied to yield equilibriatedvanadium values of 5,000; 10,000; 20,000 and 30,000 ppm. As explainedelsewhere, vanadium, as vanadium pentoxide on regenerated sorbent, mayundergo melting at regenerator temperatures and, at these metal levels,may flow across the sorbent surface, causing particle fusion andcoalescence.

For example, at 1,000 ppm vanadium on a clay sorbent, this phenomenabegins to be observed and by 10,000 ppm vanadium, clay particlecoalescence becomes a major factor in unit operation. By utilizing thesorbent of the present invention, one can now operate an upgrading unitwith sorbent in the upper ranges of vanadium levels (30,000-40,000 ppm)without vanadium deposition causing particle coalescence or excessivesintering of the sorbent structure.

                  TABLE 7                                                         ______________________________________                                        Sorbent Addition Rates For Holding Given Vanadium                             Levels On Sorbent For Feeds                                                   With Varying Vanadium Content.                                                40,000 BBL/DAY UNIT                                                           Total                                                                         Vanadium in                                                                            lbs Metal                                                            Feed PPM Day                                                                  ______________________________________                                        Level on Equilibrium Sorbent                                                  5,000   10,000    20,000    30,000  PPM                                       0.5     1.0       2.0       3.0     WT %                                      Daily Tonnage Replacement                                                     400      5200     500      250    125    82                                   200      2600     250      125    65     42                                   100      1300     125      63     32     21                                   50       650      63       32     16     10                                   25       325      32       16     8      5                                    1        13       1.25     0.63   0.32   0.21                                 ______________________________________                                    

Table 8 demonstrates the economic differential (savings in dollars perday) that can be realized by utilizing the sorbent of this invention andoperating at the 30,000 ppm level versus the 10,000 ppm level ofvanadium on sorbent. The costs of virgin sorbent as given in Notes (2)and (4) of Table 8 are based on quoted prices for clay sorbents and usedcatalysts which are commercially available at the present time. As shownin Table 8, upgrading of a feedstock having 1 ppm vanadium for FCCoperations with the sorbent of the invention would result in a savingsof about $504/day. In comparison, upgrading a heavy hydrocarbon oilcontaining 25 to 100 ppm vanadium for RCC operations utilizing thesorbent of the invention would result in a savings of about $12,900 toabout $50,400 per day, depending on the vanadium content of thefeedstock.

                  TABLE 8                                                         ______________________________________                                        Feed - 40,000 Bbl/Day; 1 Bbl = 335 Lbs                                               Sorbent Add'n Rate                                                                           Sorbent Add'n Rate                                             To Maintain 10,000                                                                           To Maintain 30,000                                      Metals  Ppm V.sup.(1) Ppm V.sup.(2)                                           In Feed,         Cost            Cost   Savings                               Ppm    Tons/Day  $/Day.sup.(2)                                                                          Ton/Day                                                                              $/Day.sup.(4)                                                                        $/Day.sup.(5)                         ______________________________________                                         1     0.63        567    0.21     63     504                                 25     16        14,400   5      1,500  12,900                                50     32        28,800   10     3,000  25,800                                100    63        56,700   21     6,300  50,400                                ______________________________________                                         .sup.(1) Based on sorbent coalescence at this level for that material         described in literature                                                       .sup.(2) Sorbent cost of 45¢/lb.                                         .sup.(3) Based on sorbent of the invention coalescing at this level           .sup.(4) Sorbent cost of 15¢/lb.                                         .sup.(5) Saving is equal to .sup.(2) minus .sup.(4)                      

What is claimed is:
 1. A process for upgrading residual oil portions ofcrude oils comprising metal contaminants and Conradson carbon producingcomponents to provide an upgraded residual oil product reduced in metalcontaminants and Conradson carbon producing components whichcomprises,(a) contacting at an elevated temperature said residualportions of crude oil with solid sorbent particle material of relativelyhigh pore volume and surface area sufficient to immobilize depositedvanadium compounds by adsorption thereof within the pore structure ofsaid sorbent particle during said contacting, (b) said sorbent particlebeing selected from one of deactivated, spent or equilibrium crackingcatalyst which has been treated with one or a combination of alkalimetal compounds in an amount sufficient to neutralize available acidcracking sites therein and yield a deactivated cracking catalyst sorbentmaterial with essentially no significant cracking activity of high porevolume and surface area, and (c) using said deactivated crackingcatalyst sorbent material of step (b) in step (a) above.
 2. The processof claim 1 wherein the amount of said added alkaline metal is sufficientto lower the MAT activity of said catalyst by at least about 25 volumepercent conversion.
 3. The process of claim 2 wherein the amount of saidadded alkaline metal is sufficient to lower said MAT activity by atleast about 35 volume percent conversion.
 4. The process of claim 1wherein the MAT activity of said sorbent is in the range of about 0 toabout 10 volume percent conversion.
 5. The process of claim 1 whereinthe sorbent material comprises catalyst particles withdrawn from an FCCand/or a RCC cracking operation of substantially reduced crackingactivity and which is deactivated catalytically with an alkaline metaladditive so as to provide a MAT relative activity of less than about1.0.
 6. The process of claim 1 wherein said vanadium compounds comprisevanadium oxides, sulfides, sulfites, sulfates or oxysulfides.
 7. Theprocess of claim 1 wherein the sorbent comprises a deactivated, spent orequilibrium catalyst withdrawn from an FCC or RCC cracking operation andtreated with an alkaline metal additive so as to have a MAT relativeactivity in the range of about 0 to about 0.1%.
 8. The process of claim1 wherein the oil feed is a reduced crude or crude oil containing 100ppm or more of metals comprised of nickel, vanadium, iron or copper andhaving a Conradson carbon value of 8 wt% or more.
 9. The process ofclaim 1 wherein the addition of the alkaline metal compound may be madeduring the contact step for upgrading the residual oil feed, bytreatment of deactivated catalyst particles and/or prior to use in saidresidual oil upgrading contact step.
 10. The process of claim 1 whereinthe sorbent particle material is of a particle size for use in a fluidor moving bed solid particle contact operation.
 11. The process of claim1 wherein said upgraded product contains 50 ppm or less of metals andless than 8 wt% Conradson carbon.
 12. The process of claim 1 whereinsaid residual oil feed contains sodium salts in a concentration in therange of about 1 ppm to about 50 ppm.
 13. The process of claim 1 whereinthe residual oil feed comprises gas oil, about 0 to about 25 weightpercent of a reduced crude, more than about 0.1 ppm vanadium and has aConradson carbon value greater than about 1.0.
 14. The process of claim1 wherein said oil feed is a reduced crude or crude oil containing 75ppm or more of vanadium and having a Conradson carbon value of 10 wt% ormore.
 15. The process of claim 1 wherein said sorbent is prepared from acatalytic material comprised of a zeolite in a matrix composition, andhas a surface area of at least 20 m² /g and a pore volume of at least0.2 cc/g.
 16. The process of claim 1 wherein said sorbent is in aspherical form and ranges in size from about 10 to about 200 microns,and wherein said upgrading zone comprises a riser transfer zone.
 17. Theprocess of claim 1 wherein said sorbent ranges in size from about 200microns to about 1/4 inch, and wherein said upgrading zone comprises amoving bed contact zone.
 18. The process of claim 1 wherein prior toaddition of said alkaline metal compound said catalyst contained from 1to 20 weight percent of a catalytically active aluminosilicate zeolite.19. The process of claim 1 wherein the upgraded product of said residualoil feed is subsequently contacted with an active conversion catalyst ina catalytic conversion process to produce gasoline boiling range productof improved octane rating.
 20. The process of claim 1 wherein said addedalkaline metal comprises Li, Na, K, Rb, Cs, Mg, Ca, Sr, or Ba.
 21. Theprocess of claim 1 wherein said alkaline metal is present in the sorbentin the range of about 0.2 to about 5 wt.%.
 22. The process of claim 1wherein prior to addition of said alkaline metal said catalyst comprisesan aluminosilicate zeolite embedded in a catalytically active matrix.23. The process of claim 1 wherein prior to addition of said alkalinemetal said catalyst comprises a catalytically active amorphoussilica-alumina material containing no zeolite.
 24. The process of claim23 wherein said catalytically active amorphous silica-alumina materialis promoted with titania, zirconia or magnesia or a combination of 2 ormore of said promoters.
 25. The process of claim 1 wherein said alkalinemetal is introduced into said catalyst by contacting said catalyst withan aqueous solution of a salt of said alkaline metal or a hydrocarbonsolution of an organo-metallic compound of said alkaline metal duringone or more steps of said upgrading process.
 26. The process of claim 1wherein said deposited metals include vanadium deposited on said sorbentin concentration ranges of about 0.05 to 5 wt% based on weight ofsorbent.
 27. The process of claim 1 wherein said oil feed contains bothnickel and vanadium and the weight ratio of said vandium to said nickelis in the range of about 1:3 to about 5:1.
 28. The process of claim 1wherein said oil feed has a significant content of heavy metals,including vanadium, and the vanadium proportion of said heavy metalscontent is greater than fifty percent.
 29. The process of claim 1wherein said alkaline metal is added to said catalyst as a water solubleinorganic alkaline metal salt comprised of a halide, nitrate, sulfate,sulfite, or carbonate or a combination of two or more of said alkalinemetal salts.
 30. The process of claim 1 wherein said alkaline metal isadded to said catalyst as a hydrocarbon soluble alkaline metal compoundcomprised of an alcoholate, ester, phenolate, naphthenate, carboxylateor dienyl sandwich compound, or a combination of two or more of saidalkaline metal compounds.
 31. The process of claim 1 wherein theresidual acidity of the catalyst is determined and said alkaline metalis added to the catalyst in an amount sufficient to give an alkalineneutralization ratio of at least about 1:1.
 32. The process of claim 1wherein the residual acidity of the catalyst is determined and saidalkaline metal is added to the catalyst in an amount sufficient to givean alkaline neutralization ratio in the range of about 1.2 to about 2.0.33. The process of claim 1 wherein said alkaline metal comprises Na, K,Mg, Ca, Ba, or a combination of two or more of said alkaline metals.